Acknowledgement

It gives me great pleasure and satisfaction to express my gratitude to my mentor Mr. Tapas Kumar Bag (HOD, Process Engineering) for overseeing my work, helping me with my project and for being a continuous source of motivation during my eight-week internship. Special thanks to Ms. Divya Pandey (Official HRD) and other HR personnels for facilitating the project. I would like to thank Mr. K.K. Tyagi (Process Engineer), Mr. Vikas (Sr. Engineer, Process Engineering), Mr. Animesh Jha (Engineer, Process Engineering) and the other IGFL staff members who helped me in their own different ways and were always responsive to my queries.

Introduction

Indo-Gulf Fertilizers Ltd. (Jagdishpur) is one of the ten gas based nitrogenous fertilizer projects, set up by the Government of India. It uses large quantities of natural gas available from Bombay High/Bassein Fields. The project was promoted by Pradeshiya Industrial and Investment Corporation U.P. (PICUP) along with Gulf Consolidated Company for Services and Industries, a multi-national company of Bahrain and later on taken over by Aditya Birla group.

This fertilizer plant was set up in Jagdishpur (Sultanpur) in January 1985. It uses natural gas as its raw material as well as fuel. It is designed to produce 1350 MT of Ammonia and 2200 MT of urea, in two streams each of 1100 MT per day. The plant started commercial production on 1st November 1988.

Project Setting

The fertilizer plant of Indo Gulf Fertilizers and Chemical Corporation Ltd. (IGFCCL) has been set up near Jagdishpur in District Sultanpur on plots of land developed by the U.P. State Industrial Development Corporation (UPSIDC) for location of Industries. UPSIDC acquired an area of 684.4 hectares (1711 acres) 294.4 hectares (736 acres) in Sultanpur District and 390 hectares (975 acres) in Rae Bareli District along the National Highway No. 56 at a distance of about 78 kms. from Lucknow. The industrial area has 313 developed plots, out of which 290 plots have come up in this industrial and other units have come up in this industrial area. IGFCCL have been allotted an area of 333.332 hectares (833.33 acres). The plant of IGFCCL its township, demonstration farm and vocational training center are situated in Sultanpur district. Title of Indo Gulf Fertilizers and Chemicals Corporation Ltd. is now changed to Indo Gulf Fertilizers Ltd. (IGFL).

PROJECT SYNOPSISProduct Urea : 726000 MT/YRNutrient output N : 338750 MT/YRCAPACITY OF MAIN PLANTSAmmonia : 1x1910 MTPDUrea : 2x1600 MTPDPROCESS LICENSEProcess License for Ammonia Plant : Haldor Topsoe, DENMARKProcess License for Urea Plant : Snamprogetti, SpA Milano, Italy

Technology

The project comprises a single stream Ammonia Plant of 1350 MTPD capacity based on Haldor Topsoe Technology (Denmark) where natural gas is used as feed and fuel. IGFL has two units Urea Plant of 1100 MTDP capacity each. Urea plant is based on Snamprogetti Ammonia Stripping Technology (Italy). Annual installed capacity of the plant was 7,68,000 MT Urea.

For Steam Generation there are two service boilers, each having a capacity of generating 100 TPH of high-pressure steam (105 Ata/515C). Power is generated by two units (1 main + 1 stand by) of Gas turbine of 18 MW each, along with common HRSG to generate medium pressure steam (39 Ata/395C).

The plant is designed on multiple systems by using NG/Naphtha/HSD for gas turbine and NG/FO/LSHS/Naphtha for boilers to overcome any unforeseen shortage of natural gas.

BRIEF PROCESS DESCRIPTION OF AMMONIA PLANT:-

Ammonia Plant Processes

Ammonia is produced from a mixture of hydrogen and nitrogen. Hydrogen is made available by the processing of natural gas with steam, whereas Nitrogen from atmospheric air.During normal operational conditions, the ammonia is sent to Urea Plant. However, two ammonia storage tanks, of 10,000 & 5000 MT capacity respectively, are also provided to handle any emergency situation. Ammonia is stored in these tanks at -33C and atmospheric pressure. The process flow sheet for ammonia production is delineated in above Fig.

Ammonia Plant is designed to produce1350 MTPD of Liquid Ammonia37073 Nm3/hr of Carbon Di Oxide

Main steps for manufacture of Ammonia is -

DesulphurisationNatural Gas (lean) mainly containing methane (98%) & higher hydrocarbon has 10-50 ppm of sulphur as Hydrogen Sulphide. In this process sulphur content is reduced to 0.05 ppm by weight. Natural Gas at a pressure of 39 kg/cm2 g & 390 0C is mixed with Hydrogen & heated up in Natural Gas Preheaters. After that it passed to Desulphurisation Reactor containing Ni-Mo catalyst. Here Sulphur forms Hydrogen Sulphide on reaction with Hydrogen. Further this Hydrogen Sulphide is absorbed on a bed of Zinc Oxide Absorber.

Ni-Mo H2 + S H2S H2S + ZnO ZnS + H2O

Primary Reformer It consists of a gas fired tubular primary reformer with flue gas heat recovery section. Natural Gas is reformed with steam to produce a mixture of Hydrogen, Carbon Monoxide, Carbon Di Oxide & Methane (11.12%). The desulphurised Natural Gas & Hydrogen mixture from Zinc Oxide bed at the pressure of 34 kg/cm2 g is mixed with excess quantity of superheated steam at a pressure of 37 kg/cm2 g & 381 0C. This mixture is preheated with flue gases of primary reformer furnace at 520 0C and sent to Primary Reformer. It consists of tubes packed with Ni catalyst and exit gas temperature from the tube is 797 0C.

Natural Gas + Steam Hydrogen + {CO + CO2 + CH4}

Secondary Reformer Function of secondary reformer is to reform the primary reformer exit residual methane & to add nitrogen to the process gas. The exit gases from the Primary Reformer are mixed with a Stoichiometric quantity of compressed preheated air at 32 kg/cm2 g & 550 0C in Secondary Reformer through air nozzle. Secondary reformer outlet gas mixture contains Hydrogen, Carbon Monoxide, Carbon Di Oxide & Methane (0.3%).

Waste Heat Recovery High temperature process gas exit Secondary Reformer is cooled down in waste heat boiler & turn high pressure steam is produced. Steam produced is utilized ass process steam & as driving force for major rotary equipment. Secondary Reformer exit gases at 988 0C enter the tube side of Waste Heat Boiler having an internal bypass for temperature control. The process gas temperature at Waste Heat Boiler exit is maintained at 360 0C. On shell side of Waste Heat Boiler high pressure steam is generated at 120 kg/cm2 g & 324 0C.

Shift Conversion Carbon Monoxide is converted to Carbon Di Oxide in two stage shift conversion bringing down the Carbon Monoxide content from 12.95% to 0.3% in the process gas. This operation is carried out at two temperature levels,

High Temperature Shift Conversion The process gas exit Waste Heat Boiler with 12.95 mole% Carbon Monoxide content enters high temperature shift converter containing Iron catalyst. Here Carbon Monoxide reacts with steam to form Carbon Di Oxide & Hydrogen. Temperature in High Temperature Shift Converter catalyst bed rises to 427 0C as the reaction is exothermic in nature. Gas exit High Temperature Shift Converter contains 3.17 mole% Carbon Monoxide.

Low Temperature Shift Conversion The gas exit High Temperature Shift Converter is cooled down to a temperature of 210 0C & heat is utilized to produce steam in High Pressure Waste Heat Boiler. Temperature inlet is maintained by cooling down the gas in B.F.W. Preheater. Converter contains Copper- Zinc as catalyst. Here the water gas shift reaction proceeds further & Carbon Monoxide content in the gas mixture is brought down to 0.30 mole%.

Carbon Di Oxide Removal

Giammarco Vitrocoke (G.V.) solution is used to absorb Carbon Di Oxide from the process gas thus reducing Carbon Di Oxide content from 17.72% to 0.1% in gas mixture. G.V. solution rich in Carbon Di Oxide is regenerated liberating Carbon Di Oxide which is sent to Urea plant.

Carbon Di Oxide Absorption

The gas leaving LP boiler at 124 C temp is ccoled down in LP BFW preheater to 108 C temp. The condensate from the process gas is separated in the separator. Process gas containing 17.69 mole% carbon dioxide enters the bottom of the absorber. Separated condensate is sent to process condensate stripping section. Major part of the GV soln from LP regenerator enters in the middle of absorber at 190 C temp. Rest of the GV soln is cooled down in DMW preheater and split stream air coolers to a temp of 60 C which enters the rising process gas. The gas comes out of absorber at 60 C temp containing 0.09 mole% carbon dioxide and enters separator OH absorber. Condensate is separated in separator and process gas from the top of separator and process gas from the top of separator is sent through heat exchangers to methanator. GV soln containing carbon dioxide at 116 C temp is sent to regenerator through hydraulic turbine for carbon dioxide stripper.

Carbon Di Oxide Stripping

GV soln from absorber enters top of HP regenerator. Major part of the soln flows down the HP regenerator where as about 20 % of the soln is drawn from draw off Pan of the HP regenerator and fed to the top of LP regenerator. Heat of regeneration in regenerator is supplied through:

1. LT shifted gas in reboilers 2. LP steam through ejectors of flash tank3. LP steam through ejector of LP boiler.Carbon Dioxide from top of regenerator at 106 C temp is cooled down in LP BFW pre heater. GV soln in LP regenerator is flashed to a pressure of 0.29 Kg/cm2g. Carbon dioxide evolved is cooled down in condenser OH LP regenerator. Condensate from the gas is separated in the separator and it is sent to the process condensate stripping section. Carbon dioxide pressure from the separator is boosted in carbon dioxide booster to a pressure of 1 Kg/cm2g and this stream meets the carbon dioxide from regenerator at down stream of LP BFW preheater HP regenerator.Combined stream of carbon dioxide gas from both the regenerators enters the OH condenser where it is cooled to a temp of 40 C by cooling water . the condensate is separated in separator OH regenerator and it is sent to process condensate stripping section. Product carbon dioxide from the top of the separator sent to urea plant at 0.88 Kg/cm2g. Partially regenerated GV soln from the regenerator enters the flash tank B-305 at 126 C temp. From the flash tank the soln flows to LP regenerator F-303 bottom for further regeneration. GV soln from LP regenerator is pumped to the absorber through circulation pumps P-301 A/B/C. Process gas at 164 C temp enters GV reboilers and is cooled down to 134 C temp by exchanging heat with GV soln. It is further cooled down to a temp of 125 C in LP boiler. Shifted gas is then sent to absorber through LP BFW pre heater.

Methanation

Oxides being harmful for synthesis converter catalyst, Carbon Monoxide & Carbon Di Oxide are converted into Methane in this step. The exit process gas (Synthesis Gas) contains mainly a mixture of Hydrogen & Nitrogen in the ratio 3:1. Effluent gas from absorber at a temperature of 60 0C is heated to a temperature of 302 0C in Gas Gas Exchanger is further heated in Trim Heater to achieve a temperature of 320 0C of the mixed gas stream at the inlet of Methanator. The gas enters Methanator where final traces of Carbon Monoxide & Carbon Di Oxide present in the process gas are converted to Methane by reacting with Hydrogen in presence of Ni catalyst. Methanation reaction being exothermic in nature the gas temperature increases to 343 0C in the catalyst bed. Gas exit Methanator is cooled down in Gas-Gas exchanger by exchanging heat with Methanator inlet process gas to 85 0C. Synthesis gas mixture containing 0.8mole% Methane flows to the Synthesis Gas Compressor at a pressure of 26.0 kg/cm2 g.

3H2 + N2 2NH3

Synthesis Gas Compression

Synthesis of Hydrogen & Nitrogen takes place at high pressure, hence the gas mixture is compressed to 220 kg/cm2 g in the synthesis gas Compressor.

Ammonia Synthesis

Conversion of this gas mixture into Ammonia takes place in TOPSOE 200 radial flow converter.

Refrigeration System

Gas mixture containing Ammonia is cooled down to separate liquid Ammonia from the gases by providing Ammonia refrigeration. Separated liquid Ammonia is sent to Urea plant at 120 C.

Material and Energy Balance across 02 Section of Ammonia Plant consisting of:

1- Primary Reformer (H-201)2- Secondary Reformer (R-203)3- HT Shift-Convertor (R-204)4- LT Shift Convertor (R-205)

EFFICIENCY OF CARBON DIOXIDE REMOVAL SECTION:Input:Total Carbon Dioxide from LT shift conveter = 48812.371 Nm3/hrTotal Carbon Monoxide from LT shift conveter = 622.237 Nm3/hr Output:Total Carbon Dioxide absorbed at F-302 = 48608.25 Nm3/hrEfficiency:- Efficiency = (total CO2 absorbed/total CO2 input) x 100 = 99.58% Therefore, efficiency of absorber is 99.58%.LOSSES:-1: As we can see absorber exit gas has 0.09 mole% carbon dioxide and 0.29 mole% carbon monoxide. As this stream goes to methanation column where CO2 and CO are reacted with hydrogen to produce methane as follows: CO + 3H2 = CH4 + H2O + heat CO2 + 4H2 = CH4 + 2H2O + heatNow seeing the above reaction we can derive that due the slip CO2 and CO hydrogen gas is being burnt, which is a loss in a plant.Total hydrogen gas used = 2789.6523 Nm3/hrPRIMARY REFORMER MATERIAL BALANCE CALCULATIONS:

Feed Analysis:1- Natural GasFlow Rate- 50300 Nm3/hrComposition:ComponentPercentage by VolumeVolumetric Flow(NM3/HR)

Methane(CH4)98.6549621

Ethane(C2H6)1.06533

Propane(C3H8)0.04321.6

N20.1420370.42

CO20.096748.6 215508

2- Process SteamFlow Rate- 129.11 Tonnes /Hr Or 162802 NM3/Hr3- Recycled H2Flow Rate- 2460 NM3/Hr

OVERALL FEED ANALYSIS-OVERALL FLOW RATE = 223166 NM3/HR

Composition-

ComponentPercentage by Volume

Methane23.02

Steam75.54

Hydrogen1.14

Rest0.32

Temperature before Introduction in Primary Reformer = 491C = 764KPressure of the combined feed when introduced in Primary Reformer = 36.37 kg/cm2Atomic Analysis of Reformer Feed ComponentPercentage by mole

Carbon 6.85

Hydrogen 71.17

Oxygen 21.96

Steam: Carbon Ratio= 3.28

Analysis of Primary Reformer Exit-

ComponentVolumetric Flow( NM3/HR)Percentage By Volume

Methane19181 6.97

Steam11364141.34

Carbon-Monoxide147265.35

Carbon-Dioxide163625.95

Hydrogen11097240.37

Total Gas flowrate = 274882 NM3/HRDry gas flowrate = 161241 nm3/hrSteam: Dry gas Ratio = 0.704

Analysis of reforming reaction in primary Reformer H-201Following reactions are taking place in the primary reformer-

CH4 + 2 H2O CO2 + 4 H2 (1)

CO2 + H2 CO + H2O (2)

CH4 reacted in (1) = 30819 NM3/HR % Conversion of CH4 = 61.6 %CO formed in Reaction (2) = 14726 NM3/Hr

Fuel Analysis

1- Natural Gas as fuel Flowrate = 17300 NM3/HrCalorific Value: 8163.316 Kcal/Sm3

2- Tail GasFlowrate = 6993 NM3/HrCalorific Value = 2500 Kcal/SM3

ComponenetPercentage By Volume

N241.81

Ar9.97

CH434.04

H258.96

3- PGR Inlet Flowrate = 2000 NM3/HrCalorific Value = 2500 Kcal/Sm3ComponentPercentage By Volume

N223.1

Ar4.03

CH412.15

NH31.76

H258.96

4- Product GasFlowrate = 1500 NM3/HrCalorific Value = 2400 Kcal/Sm3ComponentPercentage By Volume

N29.28

Ar.99

CH4.65

H289.08

Net Composition of fuel entering the primary reformer:

ComponentVolumetric Flow(NM3/Hr)Percentage By Volume

CH419606.372.5

N2351113

H23141.211.6

Ar776.92.87

Total27035.4100

Total CO2 formed on combustion of fuel gas = 20168 Nm3/HrFlow Rate of exit flue gas = 233191 Nm3/HrOxygen Content in flue gas = 3444 Nm3/Hr

Net O2 consumed in combustion of fuel gases = 41768 Nm3/HrTotal O2 in Combustion Air = 45168 Nm3/HrTotal N2 in combustion AIR = 166380 Nm3/Hr Therefore from above calculations :

Total Combustion Air entering the reformer = 215548 Nm3/Hr

ENERGY ANALYSIS OF PRIMARY REFORMER(Radiant Zone): In the energy analysis done below basic equation of thermodynamics for stream flow is used i.e-

Net Enthalpy Change of stream = Heat Supplied from outside + Shaft Work Done on system

Using thermodynamic approach , we know that ENTHALPY IS A STATAE FUNCTION. It depends only on initial and final stage .

Step-1 Calculation of enthalpy in output and input streams of the primary reformer.

Approach Used is as follows :

REACTION 1 CH4 + 2 H2O CO2 + 4 H2

HEAT OF REACTION = 39.414 kCal/gm-molNet heat released in Reaction (1) = 53.466 MKCal

Reaction (2)CO2 + H2 CO + H2O Heat of Reaction = 9.846 Kcal/ gm-mol Net heat released in reaction (2) = 6.25 MKCalTable for specific heat Capacities :ComponentSpecific Heat capacity(Cal/gm-mol/K)

CH45.34+0.0115T

H26.62+0.0008T

Steam8.22+0.00015T

O28.27+0.00027T

N26.50+0.001T

CO6.60+0.0012T

CO210.34+0.0027T

Ar4.97

From the above data enthalpy change for a component undergoing a temperature Change can be calculated by the following equation:

Change in enthalpy for the reactants going from 770K to 298K = 1- For methane= -11.92 MKCal2- For Steam= -2.8 MKCalChange in enthalpy for products going from 298K to 1054K= 1- For methane=5.15 MKCal2- For steam = 31.45 MKCal3- For H2 = 26.4 MKCal4- For CO = 3.63 MKCal5- For CO2 = 6.97 MKCal

= 118.55 MKCalCalculation of energy released from fuel gas:

Net energy released from fuel gas = Heat of combustion of fuels Heat consumed by flue gases to reach 1328k.

Calculation of heat of combustion:1- Heat of combustion for methane = 191.7 Kcal/gm-mol2- Heat of Combustion for H2 = 57.76 Kcal/gm-mol3- Heat of combustion for C2H6 = 339.2 Kcal/gm-mol4- Heat of combustion for C3H8 = 487.3 Kcal/gm-mol

Net Heat of Combustion = 179.58 MKCal

Heat Consumed in rise in temperature of flue gases =

Heat consumed = 54.9 + 16.2 + 11.43 + 1.31 19.1 - 3.42 0.27 4.52 MkCal = 56.53 MKCal

NET HEAT RELEASED FROM FLUE GAS TO PRIMARY REFORMER = 179.58-56.53 MKCNET HEAT TO PRIMARY REFORMER = 123.05 MKcalTotal number of burners = 576Heat released per burner = 123.05/576 = 20920 Kcal/Hr

Calculation of efficiency of primary reformer radiant zone :

e =

e = 118.55/ 123.05 = 96.34 %

Secondary Reformer Analysis :Inlet feed to Secondary reformer:ComponentVolumetric Flow( NM3/HR)Percentage By Volume

Methane19181 6.97

Steam11364141.34

Carbon-Monoxide147265.35

Carbon-Dioxide163625.95

Hydrogen11097240.37

Total Gas flowrate = 274882 NM3/HRDry gas flowrate = 161241 nm3/hrSteam: Dry gas Ratio = 0.704

N2 in the exit of Secondary Reformer = 57388 NM3/HrTherefore Process Air entering the secondary reformer = 72643 NM3/HrOxygen content in process air = 15255 Nm3/Hr

Combustion Zone analysis :Considering higher affinity of combustion for H2 first we will consider combustion of H2 first.O2 available for combustion = 15255 NM3/HRH2 combusted = 30510 NM3/HrHeat released in combustion = 77.62 MKCal/Hr

Reforming Zone Analysis :

CH4 + 2 H2O CO2 + 4 H2 (1)

CO2 + H2 CO + H2O (2)

CH4 reacted in (1) = 18239 NM3/HR % Conversion of CH4 = 95 %CO formed in Reaction (2) = 16810 NM3/Hr

Analysis Of gas leaving secondary Reformer :

ComponentFlowrate(NM3/Hr)Percentage By Volume

Steam12483533.7

H213653636.9

CO315368.5

CO217937 3698734.8

CH4955.650.25

Ar686.110.1

N25738815.5

ENERGY ANALYSIS OF SECONDARY REFORMER : For calculating the enthalpy balance in the secondary reformer for the reactants and products we can the process in 4 parts to calculate net enthalpy change.

Entahlpy changes for part-1 :1- N2 = -2.045 MKCal2- O2 = -3.013- CO2 =- 6.624- CO = -3.635- H2 = -26.466- H2O = -31.497- CH4 = -8.37Total = -81.625 MKCal/hrEnthalpy changes for part-2 :1- Heat evolved in combustion of H2 = -77.6 MKCal/Hr2- Heat consumed in Reaction (1) = 31.65 MKCal/Hr3- Heat consumed in Reaction (2) = 7.28 MKCal/HrTotal = -38.67 MkCal/hrEnthalpy changes for part-3:1- N2 =10.09 MKCal2- CO2 = 9.383- C0 =9.994- H2 = 41.585- H2O = 43.166- CH4 = 0.577- Ar = .06Total = 114.83 -5.465 MKCal/Hr Therefore heat lost in Secondary reformer = 5.465MKCal/Hr

HT SHIFT CONVERTOR ANALYSIS :Shift Conversion Reaction :C0 + H2O CO2 + H2 + heatHeat of reaction = -9.846 Kcal/molCO Reacted = 24003 Nm3/HrPercentage Conversion of CO= 76.11 %Shift convertor exit gas analysis:ComponentVolumetric Flowrate(NM3/Hr)Percentage by vol

H216053943.4

CH49410.25

CO75332.03

CO24194411.34

N25738715.5

H2O10082927.2

Ar699 0.18

Energy Analysis of shift convertor :Inlet temperature in shift convertor = 356 deg celciusOutlet Temperature = 426 deg celciusUsing the same thermodynamic approach as above :Heat released in reaction =- 10.63 MKCal/hrSensible heat of reactants and products :FOR UNREACTED COMPONENTS :1- CH4 =.0332- Ar = .013- CO = 0.1684- CO2 = 0.6525- N2 = 1.2456- H2O = 2.587- H2 = 2.96Sensible heat of reacting species = 2.52 MKCal/hr Net sensible heat = 10.46 MKCal/hrHeat lost in HT Shift Convertor = Heat released Net Sensible heatHeat Lost = 0.175 MKCal/hr

LT SHIFT CONVERTOR ANALYSIS :CO Reacted = 6871 NM3/hrPercentage Conversion = 91.2%Heat evolved in Reaction = 2.98 MKCal/HrExit Gas Analysis

ComponentVolumetric flowratePercentage

CO24881213.19

Steam9394325.39

CO662 3698840.17

H216743545.26

N25739315.51

CH49500.256

Ar6890.186

Energy Analysis :Heat released in exothermic reaction = 2.98 MKCal/hrSensible heats for rise in temperatures = 1- CO=0.00392- CH4=0.00863- Ar=0.00284- H2O=0.6515- H2=0.9416- N2=0.337- CO2=0.408Sensible heats of components undergoing reaction = 0.26 MKCal/hrNet Sensible Heat = 2.605Net Heat Lost = Net Heat released Net Sensible heatNet Heat Lost = 0.375 MKCal/hr

Energy Generated by steam for Compressors :

Steam Requirements :

K-421: HS: P=38 Kg/cm2 T= 380 C Flow rate: 61.94 T / hr H = 3148 KJ/ Kg Net Energy Generated = 194987120 KJ/hr

K-451 : HS: P= 38 kg/cm2 T= 380 C

Flow rate : 15.68 T/hrH=3148 KJ/hrNet Energy Generated = 49360649 KJ/hr

05-Section overall balanceFeed Analysis: Make up Gas :Flow Rate- 212000 Nm3/hr or 9464.785 Kmol/hrComposition:ComponentPercentage by mole

Methane(CH4)1.02

N226.30

Ar0.4

H272.28

Product Analysis: Product H2 :Flow rate : 11508 Nm3/hr or 513 Kmol/hrComposition:ComponentPercentage by mole

Methane (CH4)0.31

N26.33

Ar0.67

H292.69

Purge Gas :Flow rate : 1700 Nm3/hr or 75.888 Kmol/hrComposition:ComponentPercentage by mole

Methane (CH4)13.8

N222.08

Ar4.43

H252.89

NH36.8

Tail Gas :Flow rate : 6000 Nm3/hr or 267 Kmol/hrComposition:ComponentPercentage by mole

Methane (CH4)33.61

N244.78

Ar10.08

H211.53

NH3 to Urea plant: 79200 Nm3/hr or 3535 Kmol/hr NH3 to storage : 3640 Nm3/hr or162.4 Kmol/hr

Elemental Carbon Balance :Inlet : 9464.785 X 0.0102 = 96.535 Kmol/hr

Outlet : 513 X 0.0031 + 75.88 X 0.138 +267 X 0.3361 = 91.329 Kmol/hr

Carbon Loss =Inlet Outlet=96.535 91.329=5.206 Kmol/hr or 62.472 Kg/hr

Elemental Hydrogen Balance :Inlet :9464.785 X 2 X 0.0102 + 9464.785 X 0.7228 =6937.687 Kmol/hr

Outlet : 513 X 2 X 0.0031 + 0.9269 X 513 + 75.888 X 2 X 0.138 + 75.888 X 1.5 X0.068 + 267 X 2 X 0.336 + 267 X 0.1153 + 3535 X 1.5 + 162.4 X 1.5=6235.042 Kmol/hr

Hydrogen Loss = Inlet Outlet=6937.687 6235.042=702.63 Kmol/hr or 15738 Nm3/hr or 1405.276 Kg/hr

Latest Technologies for Enhancing the Capacity and Fuel Consumption of Ammonia PlantProposal 1:

KBR Reforming Exchanger System

KRES is a proprietary heat exchanger-based steam reforming technology consisting of a fired preheater, an autothermal reformer (ATR) and a reforming exchanger. KRES takes the place of a conventional primary reformer by feeding excess air, natural gas feed and steam to the ATR and feed and steam in parallel into the upper end of the robust, shell-and-tube reforming exchanger. The compact ATR and reforming exchanger in combination with the fired preheater take up much less plot space than a conventional fired steam methane reformer.

The tubes in the KBR reforming exchanger are open-ended and hang from a single tube sheet at the inlet cold end to minimize expansion problems. They are packed with a conventional reforming catalyst, which can be easily loaded through a removable top head. The tubes are accessible and removable as a bundle for maintenance. This simple, proprietary design has proven to be extremely reliable and maintenance free in commercial operation since 1994.

Heat to drive the reforming reaction is supplied by the effluent gas from the ATR, which operates in parallel with the reforming exchanger. To ensure adequate heat to drive the reaction, the ATR receives excess process air, typically 50 percent more than what is required for nitrogen balance.The hot ATR effluent enters the lower shell side of the reforming exchanger where it combines with reformed gas exiting the reforming tubes. This combined gas stream travels upward through the baffled shell side of the reforming exchanger providing heat needed for the endothermic reforming reactionoccurring inside the catalyst-filled reforming tubes. In this way, heat energy that would otherwise be used to generate possibly unneeded steam in a waste heat boiler downstream of the reformer is used, instead, to replace fuel as the source of heat to drive the reforming reaction.

Process Flow Diagram of KRES Scheme:

Revamping a plant for capacity increase using conventional technology has several limitations as the throughputs increase in various sections of the plant. Examples are: The size of the reformer furnaces radiant and convection sections needs tobe increased. Further, the capacity of the flue gas side also needs to beincreased. Pressure drop through the furnace is increased, which reduces the front-endpressure, including the syngas machine suction pressure. This wouldtypically reduce plant energy-efficiency and will make revamping of thesyngas machine more expensive. Heat duty of the reformed gas waste heat boiler increases, making operationof this critical equipment more severe. Typically, equipment associatedwith the steam system waste heat boiler, steam drum, steam superheater,boiler feed water pumps, piping may all need modification/change whichis expensive.

KRES technology makes revamping of an ammonia plant more economically attractive by resolving the above issues as follows:

The reforming exchanger adds up to 25% capacity in the reforming sectionas additional feed gas and steam mixture is processed in it by recovering thehigh-grade waste heat available exit the secondary reformer. Duties on thefurnace services are either maintained or increased only as far as the existingcoils/tubes can take it. As confirmed during the engineering phase ofseveral revamp projects, the new small mixed-feed coil for the reformingexchanger is easily accommodated in the existing furnace system. As the convection duty of the existing furnace is increased relative to itsradiant duty (post revamp compared to the base case), furnace efficiency isimproved in the revamp case. As the reforming exchanger handles the additional throughput in parallel tothe existing equipment without significantly increasing flows through theexisting key equipment having relatively high pressure drop, the issue offront-end pressure drop has a resolution. The reformed gas waste heat boiler now receives a cooler feed gas at 700 to850C which makes conditions for this critical equipment less severe.Where applicable, the concern of potential waste heat boiler tube failure canbe addressed by providing such mild conditions, and hence, such a criticalequipment need not be modified or replaced. As changes in the steam system flow rates can be minimized, replacementor modification of the high-pressure steam system (e.g. steam drum, wasteheat boiler, steam superheater, BFW pump and piping) can be avoided orminimized, which contributes to project viability.

Proposal 2:

KBRs Purifier Ammonia Process :

KBRs Purifier Ammonia Process combines the following proprietary technologies to yield an extremelyreliable, robust, low-energy plant: Mild reforming with excess air KBR Purifier Magnetite ammonia synthesis in a horizontal converterWith KBRs cryogenic Purifier syngas technology, you receive a lower-cost, more robust processing route to high-purity synthesis gas in ammonia manufacturing plants. The proprietary, front-end process technology simultaneously removes impurities (i.e., methane, argon) from synthesis gas by washing it with excess nitrogen while adjusting the hydrogen to nitrogen (H2/N2) ratio to 3:1.Benefits of Purifier Ammonia Process

Providing a clean, dry, make-up gas to the synthesis loop and simple and precise H2/N2 ratio control, Purifier technology offers benefits to the entire operation in several key areas:Low Energy Consumption A clean, dry make-up gas reduces the load on the synloop compressor and refrigeration systems, providing operational cost savings Mild reforming temperatures are used as methane slip is unimportant, which reduces fuel consumption and increases tube life Higher loop conversion is achieved with low inerts Purifier plants operate at the lowest proven energy consumption; a recent plant achieved an energy consumption of 6.5 Gcal/MT(ISBL, LHV basis)Purifier Ammonia Pro Reduced Capital Costs No separate purge gas recovery unit is needed because purge gas rejected from the synloop is passed through the Purifier unit Very clean make-up gas provided by KBRs Purifier lowers synthesis pressure, catalyst volume and purge rate, which means that smaller synloop equipment can be used

Flexibility Achieves greater stability and flexibility of operation, since the reforming section does not need tobe tightly controlled to produce a precise H2/N2 ratio Maintains production even in the event of catalyst deactivation upstream of the Purifier

Reliability Low reforming temperatures translate to lower stresses in and longer life of reformer tubes Numerous Purifier plants have run 3 - 4 years without a maintenance sh

PROPOSAL 3:Objective: To increase the absorption of CO2 over absorber F-302, thus increasing the efficiency of absorber and reducing the hydrogen consumption at methanator R-311. To achieve the above objective we should replace MEA( Mono Ethanol Amine) solution with MDEA( Methyl DiEthanol Amine) solution. Advantages of MDEA Process:1: MDEA is more efficient carbon dioxide absorber than MEA solution. Results show that MDEA solution is able to absorb 99.95 % CO2 compared to 99.58 % CO2 absorbed by MEA solution. In return less hydrogen gas consumption and less energy wastage.2: Low inerts make-up in synthesis gas due to low CO2 slip. Low CO2 slip with product gas has further advantages such as less consumption of H2 in methanator, higher conversion per pass in synthesis converter due to less inerts and reduction in purge gas from synthesis loop.3: Utilisation of all the existing equipments i.e. not making any changes in the system.4: Lower MDEA make up requirement. 5: MDEA (Methyl diethanol amine) is environment friendly and biodegradable chemical. MEA solution is a corrosive solution, while MDEA is non-corrosive. Hence MDEA system does not require any corrosion inhibitor as compared to MEA solution which uses V2O5 which is carcinogenic in nature.

PROPOSAL 4:Objective: To reduce the energy consumption in Carbon Dioxide removal section and make the process more energy efficient. Existing CO2 removal section Modified CO2 removal section

If the process stream from the top of F-303 is again diverted to F-301 regenerator, and the residue of B-301 is brought back to the absorber. This modification done with MDEA solution reduces the energy requirements of the plant by 50 60% of the current energy consumption.PROFIT:By using the above combination of MDEA solution and Two stage removal process we can optimize the plant for more production. As MDEA solution absorbs 0.37% more than the existing MEA and GV solution, and preventing the use carcinogenic corrosion inhibitor V2O5, therefore cost effective.Second, by using the Two stage removal process with MDEA solution the removal process will be optimized for less energy consumption, thus less energy and more cost effective.

Following proposals are based on KAAP (Kelloggs Advanced Ammonia Process) Technology based on some advancement over traditional ammonia process technology :

Proposal 5:

Gas Heated ReformersThis technologiesincludes the use of Gas Heated Reformers (GHR), which are tubular gas-gas exchangers. In the GHR, the secondary reformer outlet gases supply the reforming heat. Though it is not presently being used widely, GHR has certain advantages over fired furnaces. Table 1 shows a list of these advantages. Although GHR results in reduced energy consumption, a comprehensive energy conservation network should be established to maximize the benefits of a GHR system.

Table 1: Advantages of Gas Heated ReformersFired FurnaceGas Heated Reformers

Large volumesSmaller volumes

Larger surface area and heat lossReduced surface area and heat loss

Complex instrumentationSimplified instrumentation

High maintenance costsLow maintenance costs

Large convection zoneNo convection zone

Stack lossesNo stack losses

High fixed capital costsLow fixed capital costs

Reduced catalyst tube loss from high temperature and uneven heat distributionLonger tube life due to uniform heat distribution

Increased downtime required for shut downReduced downtime required for shut down

Well established processYet to gain wide acceptance

Proposal 6:Hydrogen SeparationLechatelier's Principle states that a reaction equilibrium can be shifted by applying external forces. This offers a means of removing products from the reaction mixture to increase the conversion per pass. In reforming, experiments have been performed up to 5000C (9320F) and 20 bar (294 psig) using a palladium membrane to remove the product hydrogen. These experiments have results in a significant increase in methane conversion as can be seen by the following case study.Case Study on the Membrane Separation Process The separation of hydrogen from the product gas of the reforming process can result in significant productivity gains when compared to the current processes being employed. The base case for this study is a conventional steam reforming plant based on natural gas operating at 1750 tonnes per day. The operating conditions of the plant are assumed to be the same as those typically employed today and the only modification is the introduction of hydrogen separation. The tests for the membrane separation have been carried out at 5000C (9320F) and 20 bar (294 psig), these conditions will function as upper limits for the process to be considered in this study. Membrane units will be considered after the primary reformer (at 60% hydrogen separation), after the secondary reformer (at 60% hydrogen separation), and after the High Temperature Shift (HTS) converters (at a 50% hydrogen separation) The following assumptions are made in this case study:1. The natural gas feed at the primary reformer is the same for both cases.2. The primary reformer exit temperature is the same for both units.3. The primary reformer operating pressure is the same for both units.4. The process air is fed to the secondary reformer at optimal conditions and any remaining nitrogen that is required is supplied through an Air Separation Unit (ASU) and is available at 0.1 kg/cm2(1.42 psig)5. Any extra energy consumption in the ASU is considered for the revamp case.6. All of the heat from the process gas from the primary reformer to the carbon dioxide removal section is used in a steam network.7. No changes in the carbon dioxide removal system are considered.8. The pressure drop across the front end of the process is kept constant for both systems, thus the synthesis gas compressor suction pressure remains constant.9. The loop pressure is the same for both processes and is controlled by changing the purge gas quantity.10. The existing compressors are capable of handling any additional loads.11. No scheme changes are considered in the synthesis loop.12. All hydrogen from the membrane separation unit is available at 9.0 kg/cm2(128 psig)13. The productivity analysis is carried out on the ammonia plant only (the urea plant is excluded)14. A complete steam balance is included on both processes. Changes in the steam balance are considered for:Steam generation from the front end of the processesSteam generation from the back end of the processesAdditional steam in the carbon dioxide removal section caused by a reduction in the heat available from the process gasAdditional power for the synthesis compressor due to changes in flow and compositionAdditional power in the ammonia refrigeration compressorReduced load on the process air compressorAdditional power for low pressure hydrogen separated through membranesAdditional power for nitrogen compressionAdditional power for the air compressors of the ASUSmall changes in other drives and small equipment

Comparison Between Conventional Reforming and Reforming with Hydrogen SeparationProduction rise from 1750 to 1854 tonnes per day+6.0% rise in capacity

Process ChangeEnergy Change (Gcal/tonne)

Gain in feed and fuel including steam superheater+0.36

Loss in steam generation (front end)0.00

Loss in steam generation (back end)-0.02

Loss in additional steam for carbon dioxide removal-0.27

Gain in energy in synthesis gas compressor+0.01

Extra energy in refrigeration compressor0.00

Gain in energy in process air compressor+0.16

Extra power in hydrogen compressor-0.22

Extra power for nitrogen from ASU-0.12

Steam savings in primary reformer+0.08

Other rotary drives and equipment+0.04

Total Gain+0.02

It is evident from these results that the major losses occur in the carbon dioxide removal section of the plant. These losses are the result of consuming additional steam and compression energy for hydrogen separation. With additional minimization of these losses, additional savings can result. For a production gain of6%over the base case, the energy saving is0.02 Gcal/tonne (0.08 MBtu/tonne).

This development could yield savings by increasing methane conversion in reformers and increasing the carbon monoxide conversion in shift reactors. The energy savings can be as high as 0.50 Gcal/tonne (1.98 MBtu/tonne) depending on the adopted process configuration. Hydrogen separation technology can also result in increased ammonia plant capacity as illustrated in the above case study.The reduced air requirement (about 80% of conventional plants) in the secondary reformer is required with a 60% hydrogen removal rate in the reformer. This will also require an additional source of nitrogen. Therefore, the technologies in which nitrogen is being added separately, either from an Air Separation Unit (ASU) or from any other sources, will become more important in this case.

Proposal 7:Synthesis CatalystAfter almost 90 years of a monopoly in the ammonia synthesis market, iron catalyst has not been replaced by a precious metal (ruthenium) based catalyst used in the KAAP developed by Kellogg. The KAAP catalyst is reported to be 40% more active than iron catalysts.Research work on low temperature and low pressure catalysts to produce ammonia at 20-40 kg/cm2g and 1000C is being performed atProject and Development India Ltd.(PDIL) according to their in-house magazine. The catalyst is based on cobalt and ruthenium metals and has exhibited few encouraging results.