Ammonia Production Process

.4.1  Ammonia Production Process

Ammonia is produced basically from water, air and energy. The energy source is usually hydrocarbon that provides hydrogen for fixing the nitrogen. The other energy input required is steam and power. This can be through coal or petroleum products or purchased power from a utility company.

Steam reformation process of light hydrocarbon particularly Natural Gas (NG) is the most efficient route for the production of ammonia. The other routes are the partial oxidation of heavy oils if the available feedstock is residual heavy oil from a refinery. Coal has also been used to produce ammonia. The following is an approximate comparison of the energy consumption, cost of production and the capital cost of the plants for three the feedstocks.

                                          Natural Gas              Heavy Oil           Coal

Energy consumption                1.0                          1.3                          1.7

Investment cost                       1.0                          1.4                          2.4

Production cost                       1.0                          1.2                          1.7

Natural gas is therefore the most appropriate source of feedstock on all the three accounts.

Based on the known resources of fossil raw materials and economy of use on all accounts, it is likely that natural gas will dominate as feedstock for ammonia production in the foreseeable future. Coal may become a competing feedstock if the prices of natural gas and petroleum products go very high due to depleting resources.

For the present time and near future, the steam/air reforming concept based on natural gas is considered to be the most dominating and best available technique for production of ammonia. The reforming process can be divided in to the following types:

4.4.1.1   Conventional steam reforming with fired primary reformer and stoichiometric air secondary reforming (stoichiometric H/N- ratio)

4.4.1.2  Steam reforming with mild conditions in fired primary reformer and excess air in secondary reformer (Under-stoichiometric H/N ratio)

4.4.1.3  Heat exchange auto thermal reforming, with a process gas heated steam reformer (heat exchange reformer) and a separate secondary reformer, or in a combined auto thermal reformer using excess or enriched air (under- stoichiometric or stoichiometric H/N-ratio)

All the three reforming versions are in use but the conventional one is the oldest and most in use.

4.4.1.4  Conventional Steam Reforming:

Overall conversion

The theoretical process conversions, based on methane feedstock, are given in the following approximate formulae:

0.88CH₄ + 1.26 Air + 1.24 H₂O ¾®  0.88CO₂ + N₂ + 3H₂

N₂ + 3H₂¾® 2NH₃

The synthesis gas production and purification normally takes place at 25 to 35 kg/cm² pressure. The ammonia synthesis pressure is in the range of 100-250 kg/cm². The block diagram of the steam/ air reforming is as under (Figure 4.4.1.4).

P

Feedstock desulphurisation

This part of the process is to remove the sulphur from the feedstock over a Zinc oxide catalyst-bed, as sulphur is poison to the catalysts used in the subsequent processed. The sulphur level is reduced to less than 0.1 ppm in this part of the process.

Primary reforming

The gas from the desulphuriser is mixed with process steam, usually coming from an extraction turbine, and steam gas mixture is then heated further to 500-600° C in the convection section before entering the primary reformer. In some new or revamped plants the preheated steam/gas mixture is passed through an adiabatic pre-reformer and reheated in the convection section before entering the primary reformer.

The amount of process steam is given to adjust steam to carbon-molar ratio (S/C- ratio), which should be around 3.0 for the reforming processes. The optimum ratio depends on several factors, such as feedstock quality, purge gas recovery, primary reformer capacity, shift operation and the plant steam balance. In new plants, S/C ratio may be less than 3.0.

The primary reformer consists of a large number of high-nickel chromium alloy tubes filled with nickel-containing reforming catalyst in a big chamber (Radiant box) with burners to provide heat. The overall reaction is highly endothermic and additional heat is provided by burning of gas in burners provided for the purpose, to raise the temperature to 780-830°C at the reformer outlet.

The composition of gas leaving the reformer is given by close approach to the following chemical equilibrium:

CH₄ + H₂O  ¬¾®    CO + 3H₂

CO + H₂O  ¬¾®     CO₂ + H₂

The heat for the primary reforming is supplied by burning natural gas or other gaseous fuels, in the burners of a radiant box containing catalyst filled tubes.

The flue gas leaving the radiant box has temperature in excess of 900°C, after supplying the high level heat to the reforming process. About 50-60% of fuel’s heat value is directly used in the process itself. The heat content (waste heat) of the flue-gas is recovered in the reformer convection section, for various process and steam duties. The fuel energy required in the conventional reforming process is 40-50% of the process feed energy.

The flue-gas leaving the convection section at 100-200° C is one of the main sources of emissions from the plant. These emissions are mainly CO₂, NOx, with small amounts of SO₂ and CO.

Secondary reforming:

Only 30-40% of the hydrocarbon feed is reformed in the primary reformer because of the chemical equilibrium at the actual operating conditions. The temperature must be raised to increase the conversion. This is done in the secondary reformer by internal combustion of part of the gas with process air, which also provides the nitrogen for the final synthesis gas. In the conventional reforming process the degree of primary reforming is adjusted so that the air supplied to the secondary reformer meets both the heat and the stoichiometric synthesis gas requirement.

The process air is compressed to the reforming pressure and heated further in the primary reformer convection section to about 600°C. The process gas is mixed with the air in a burner and then passed over a nickel-containing secondary reformer catalyst. The reformer outlet temperature is around 1000°C, and up to 99% of the hydrocarbon feed (to primary reformer) is converted, giving a residual methane content of 0.2-0.3 (dry gas bases) in the process gas leaving the secondary reformer.

The process gas is cooled to 350-400°C in a waste heat boiler or waste heat boiler/super heater down stream from the secondary reformer.

Shift conversion: 

The process gas from the secondary reformer contains 12-15% CO (dry gas bases) and most of the CO is converted in the shift section according to the reaction:

CO + H₂O  ¬¾®    CO₂+ H₂

In the high temperature shift conversion (HTS), the gas is passed through a bed of iron oxide/Chromium oxide catalyst at around 400°C, where the CO content is reduced to about 3% (dry gas bases), limited by the shift equilibrium at the actual operating temperature. There is tendency to use copper containing catalyst to increase conversion. The gas from the HTS is cooled and passed through the low temperature shift (LTS) converter.

The LTS is filled with a copper oxide/Zinc oxide-based catalyst and operates at about 200-220° C. The residual CO content is important for the efficiency of the process.  Therefore, efficiency of shift step in obtaining the highest shift conversion is very important.

CO₂ Removal    

The process gas from the low temperature shift converter contains mainly H₂, N₂, CO₂, and excess process steam. The gas is cooled and most of the excess steam is condensed before it enters the CO₂ removal section. This condensate usually contains 1500-2000 ppm of ammonia, 800-1200 ppm of methanol and minor concentration of other chemicals. All these are stripped and in the best practices the condensate is recycled. The heat released during cooling/condensation is used for:

§  Regeneration of CO₂ scrubbing solution

§  Driving the absorption refrigeration units

§  Boiler water preheat.

The amount of heat released depends on the process steam to carbon ratio. If all this low level heat is used for CO₂ removal or absorption refrigeration, high-level heat has can be used for feed water system. An energy-efficient process should therefore have a CO₂ removal system with low heat demand.

The CO₂ is removed in a chemical or physical absorption process. The solvents used in chemical absorption process are mainly aqueous amine solutions Mono Ethanolamine (MEA), activated Methyl DiEthanolamines (aMDEA) or hot potassium carbonate solutions. Physical solvents are glycol dimethylethers (Selexol), propylene carbonates and others.

Benfield process, Selexol, aMDEA or similar processes are considered as best practice.

Residual CO₂ content are usually in the range 100-1000 ppmv, depending on the process used. Contents of CO₂ down to 50 ppmv are achievable.

Methanation

The small residual amount of CO and CO₂ in the synthesis gas, are poisonous for the ammonia synthesis catalyst and must be removed by conversion to CH4 in the methanator:

CO + 3H₂   ¾¾®  CH₄ + H₂O

CO₂ + 4H₂  ¾¾®   CH₄ + 2H₂O

The reaction takes place at around 300°C in a reactor filled with nickel containing catalyst. Methane is an inert gas but water must be removed before entering converter.

Synthesis gas compression and ammonia Synthesis

Modern ammonia plants use centrifugal compressors for synthesis gas compression, usually driven by steam turbines, with steam being produced within the ammonia plant from exothermic heat of reactions. The refrigeration compressor, needed for condensation of product ammonia, is also driven by a steam turbine.

The synthesis of ammonia takes place on an iron catalyst at pressure usually in the range of 100-250 kg/cm² and temperatures in the range of 350-550°C:

N₂ + 3H₂ ¬¾¾®   2NH₃

Only 20-30% of synthesis gas is converted to ammonia per pass in multibed catalyst filled the converter due to the unfavorable equilibrium conditions. The ammonia that is formed is separated from the product gas mixture by cooling/ condensation, and the unreacted gas is recycled with the addition of fresh make up synthesis gas, thus maintaining the loop pressure. In addition, extensive heat exchange is required due to exothermic reaction and large temperature range in the loop.

A newly developed ammonia synthesis catalyst containing ruthenium on a graphite support has a much higher activity per unit of volume and has the potential to increase conversion and lower operating pressure. This has the potential to reduce energy consumption.

Synthesis loop arrangement differ with respect to the points in the loop at which the make-up gas is delivered and the ammonia and purge gas are taken out.

Conventional reforming with methanation as the final purification step, produces a synthesis gas contains inerts (Methane and argon) in quantities that don’t dissolve in the condensed ammonia. The major part of these is removed by taking out a purge stream from the loop. The size of this purge stream controls the level of inerts in the loop to about 10-15%. The purge gas is scrubbed with water to remove ammonia before being used as fuel or before being sent to hydrogen recovery unit.

Ammonia condensation is far from complete if cooling is with water or air and is usually not satisfactory. Vaporizing ammonia is used as a refrigerant in most ammonia plants, to achieve sufficiently low ammonia concentration in the recycled gas. The ammonia vapours are liquefied by compression in the refrigeration compressor.

          4.4.2  Steam reforming with excess air secondary reforming

This process is divergent than the conventional process broadly in the following ways:

§  Decreased firing in primary reformer

§  Increased process air flow to the secondary reforming

§  Cryogenic final purification after methanation

§  Lower inert level of the make-up syngas.

In this process part of load of primary reformer is shifted to a thermodynamically more efficient secondary reformer.  However, excess nitrogen has to be removed in the gas purification step.

                      4.4.3  Heat exchange auto thermal reforming:

From thermodynamic point of view, it is wasteful to use the high-level heat of secondary reformer outlet gas and the primary reformer flue-gas, both at temperatures around 1000°C, simply to raise steam. Recent developments are to recycle this heat to the process itself, by using the heat content of the secondary reformed gas in a newly developed primary reformer (gas heated reformer, heat exchange reformer), thus eliminating the fired furnace. Surplus air or oxygen-enriched air is required in the secondary reformer to meet the heat balance in this auto thermal concept.

The developers of this technology claim better performance on energy and are trying to perfect the systems.

          4.4.4  Best available techniques (BAT) reforming process for new plants:

The modern versions of the conventional steam reforming and excess air reforming processes will still be used for new plants for many years to come. Developments are expected to go in the following directions:

i.                   Lowering the steam carbon ratio

ii.                 Shifting duty from primary to secondary reformer

iii.               Improved final purification

iv.                Improved synthesis loop efficiency

v.                  Improved power energy system

vi.                Low NO_x burners

vii.              Non iron based ammonia synthesis catalyst

In India almost all NG based plants and naphtha based plants are based on conventional steam reforming process. Some newer plants have introduced adiabatic pre-reforming, operating at low steam carbon ratio, introduced purge gas recovery to control inerts efficiently, provided low NO_x burners and improved steam & power system resulting in better performance.

          4.4.5  Partial oxidation of heavy oils

The partial oxidation process is used for the gasification of heavy feedstock such as residual oils and coal. Extremely viscous hydrocarbons may also be used as fraction of the feed.

An air separation unit is required for the production of oxygen for partial oxidation step. The nitrogen is added in the liquid nitrogen wash to remove impurities from the synthesis gas and to get the required hydrogen/nitrogen ratio in the synthesis gas.

The partial oxidation is a non-catalytic process, taking place at high pressure (>50 kg/cm2) and temperatures around 1400°C. Some steam is added for temperature moderation. The simplified reaction pattern is:

-CHn - + 0.5 O₂ ¾¾®  CO + n/2H₂

Carbon dioxide, methane and some soot are formed in addition. The sulphur compounds in the feed are converted to hydrogen sulfide. Mineral compounds in the feed are transformed in to specific ashes. The process gas is freed from solids by water scrubbing after waste heat recovery and the soot is recycled to feed. The ash compounds are drained with the process condensate and/or together with the soot. The hydrogen sulphide in the process is separated in a selective absorption step and reprocessed to elemental sulphur in a Claus unit.

The shift conversion usually has two temperature shift catalyst beds with intermediate cooling. Steam for shift conversion is supplied partially by a cooler-saturator system and partially by steam injection.

CO₂ removed by using an absorption agent, which might be the same as in the sulphur removal step. Residual traces of absorption agent and CO₂ are then removed from the process gas, before final purification by a liquid nitrogen wash. In this unit practically all the impurities are removed and nitrogen is added to give the stoichiometric hydrogen to nitrogen ratio.

Ammonia synthesis is quite similar to steam reformation plants, but more efficient due to high purity of synthesis gas from liquid nitrogen wash unit and the loop does not require a purge.

The process block diagram is as under.

Figure 4.4.5  Block Diagram Of The Partial Oxidation Process

In India presently four plants set up in 70’s are working using the partial oxidation process to use Fuel Oil or LSHS feed stocks. Due to higher energy consumption in these plants and due to higher basic cost of feedstock in comparison to NG, these would changeover to NG as feedstock

          4.4.6  Description Of Urea Production Processes

The commercial synthesis of urea involves the combination of ammonia and carbon dioxide at high pressure to form ammonium carbamate, which is subsequently dehydrated by the application of heat to form urea and water.

 2NH₃  +  CO₂   ¬® NH₂COONH₄  ¬®  CO(NH₂)₂ + H₂O

Ammonia  Carbon                  Ammonium                          Urea              Water

                    Dioxide                   Carbamate

First reaction is fast and exothermic and essentially goes to complete under the reaction conditions used industrially. Subsequent reaction is slower and endothermic and does not go to completion. The conversion (on a CO₂ basis) is usually in the order of 50-80%. The conversion increases with increasing temperature and NH₃/CO₂ ratio and decreases with increasing H₂O/CO₂ ratio.

The design of commercial processes involves three major considerations:

§  to separate the urea from other constituents,

§  to recover excess NH₃ and

§  decompose the carbamate for recycle.

The simplest way to decompose carbamate to CO₂ and NH₃ requires the reactor effluent to be depressurized and heated. Since it is essential to recover all the gases for recycle to the synthesis to optimize raw material utilization and since re-compression was too expensive an alternative was developed. This involved cooling the gases and re-combine them to form carbamate liquor, which was pumped back to the synthesis. A series of loops involving carbamate decomposers at progressively lower pressure and carbamate condensers were used. This was known as the “Total recycle process”. A basic consequence of recycling the gases was that the NH₃/CO₂ molar ratio in the reactor increased thereby increasing the urea yield.

Significant improvements were subsequently achieved by decomposing the carbamate in the reactor effluent without reducing the system pressure. This “Stripping Process” dominated synthesis technology and provided capital/energy savings. Two commercial stripping systems were developed, one using CO₂, and other using NH3 as the stripping gases.

Since the patents on stripping technology have expired, other processes have emerged which combine the best features of Total Recycle and Stripping Technologies.

The urea solution arising from the synthesis /recycle stages of the process is subsequently concentrated to a urea melt for conversion to solid prilled or granular product.

Improvements in process technology have concentrated on reducing production costs and minimizing the environmental impact. These include boosting CO₂ conversion efficiency, increasing heat recovery, reducing utilities consumption and recovering residual NH3 and urea from plant effluents. Simultaneously the size limitation of prills and concern about the prill tower off gases effluent were responsible for increased interest in melt granulation processes and prill tower emission abatement. Some or all these improvements have been used in updating existing plants and some plants have added computerized systems for process control, New urea installations vary in size from 800 to 2000 tonnes per day.

Modern processes have very similar energy requirements and very high material efficiency. There are some differences in the details of energy balances but they are deemed to be minor in effect.

Block diagram for CO₂ and NH₃ stripping total recycle processes are as shown in Figure 4.4.6a and 4.4.6b respectively.   

A list of Ammonia-Urea plants with feedstock & technology is given in Annexure 4.I.

              4.7  Factors contributing for higher energy consumption

Most of the energy consumed in fertiliser industry is in the production of nitrogenous fertiliser and that too in the production of ammonia. In the past the energy consumption per unit production has been high. The new plants have been performing much better and the energy consumption is comparable to the best in the world. The old plants have also improved their performance but have the limitation of old technology and inefficient feedstock.

The energy consumption for the production of ammonia in a modern steam reforming plant is 50-60% above the thermodynamic minimum. More than half the excess consumption is due to compression losses and release of low-level energy that is not economical to recover. The practical minimum consumption is assumed to be about 140% of the theoretical minimum.

The record of Indian fertiliser industry on energy front in the 70’s and 80’s was not been very good. There have been many reasons for the high-energy consumption. These have been analyzed as under.

          4.7.1  Low capacity utilization

The ammonia process is a continuous operation, consisting of many sub-processes, leading to the final production of ammonia. During startup lot of energy is consumed to bring the operation parameters of all the sub-processes to those levels required for operational performance & stability. Since a large ammonia plant handles large quantities of inflammable fluids, a number of safety features are built in to the processes to trip the plant and bring in to safe condition in case of a disturbance endangering the plant. This is also necessary to avoid any major accident. If for any reason any one sub-process in the ammonia production gets disturbed and the plant process goes in to dangerous operational zone, the safety system automatically actuates and the plant gets shut down.  The material in the process gets discharged in to atmosphere and burnt. Frequent shutdowns thus result in to wastage of energy. Unfortunately, the plant outages/trips have been very frequent in the 70’s, 80’s and in some plants even in the 90’s. This is indicated by the low capacity utilization of nitrogenous plants (Table 4.7.1).

Table 4.7.1  Capacity Utilization (CU) of Nitrogenous Fertiliser Plants

Year	82-83	84-85	86-87	88-89	90-91	92-93	94-95	96-97	98-99
CU (%)	67	74	79	85.2	85.7	88.1	91	93.2	99.2

Main factors for low capacity utilization are as follows:

4.7.1.1  Power Supply

The power supply from utilities was not stable causing the plant to trip due to frequent interruptions in power supply and fluctuations in voltage. Due to sensitive nature of plants trip systems are in-built to take the plant to safe condition after tripping and venting all the gases in process and burning them off. Because of this perennial problem faced by most of the plants the Government allowed each fertiliser plant to have its own captive power plants.

4.7.1.2  Steam Supply

The plants of 70’s and 80’s vintage had their steam supply from steam generation plants using coal. The quality of coal supplied to these plants has been of poor quality with very high ash content resulting in to extensive wear & tear in boilers, breakdowns and interruptions in steam supply. There were quite a few interruptions for non-availability of coal at pithead or non-availability of railway wagons to transport the coal. It also increased energy consumption.

4.7.1.3 Indigenisation of Spares

Due to non-availability of foreign exchange attempts were made to utilize spares from indigenous sources that were not proven in quality. Further because most of the plants were in Public Sector, the purchases were made from the lowest cost suppliers rather than suppliers of proven performance. 

4.7.1.4  Unreliable Instrumentation

Internationally the capacity utilizations were low as manufacturers were yet developing very high reliability machinery and process control instruments that relied largely on human factors. It is only in late 80’s that electronically controlled instruments for better/auto control and analysis was installed. With mechanical instruments many trips were caused by the mal-functioning of the instruments themselves. Besides after the plant tripped, there were no clues as to what caused it. Restart without diagnosis and corrective action would interrupt the process again with consequent lot of energy waste.

Adverse industrial relation scenario was also contributed to bad performance. The labour unions were very strong and non-cooperative during the period. Besides their level of skills was low. Despite training centers attached with each fertiliser plant the quality of manpower could not be developed fast enough as the management’s did not see the need to revise the curriculum to meet the current and future needs.

          4.7.2  Selection Of Equipment / Available Technology

The technology selection and equipment selection for the plants being set up in 70’s was not up to the mark. Besides the Indian design and consultancy organizations involved were on the learning curve. The process suppliers did not part with the best technology, sent raw hands to our detailed engineering consultants and recommended purchase of spares with original equipment that were really not needed.

Foreign exchange availability was a major limitation during the 70’s and 80’s with the result that the country had to select the process supplier who would also provide project loans. The process suppliers were further tied up with equipment manufacturers for supply of equipment with deferred payment terms. In the deal they would sell the equipment that was not proven. A number of critical equipments were supplied that resulted in to major plant limitations. The boiler feed water pumps and untried centrifugal compressors are only few examples.

          4.7.3  Feedstock

The best feedstock for nitrogenous fertiliser is NG. During the period there was urgent need to produce indigenous fertiliser with the available feedstock. The naphtha based and fuel oil based had to be put up though they were not the best feedstock with inherent high energy consumption.

Cooling water is one process material that passes through a lot of equipment for cooling. This water needs to be treated to control corrosion in the process equipment and needs proven technology and material inputs to make it suitable. Due to non-availability of foreign exchange a number of fertiliser plants experimented with un-proven technology and chemicals and the equipment suffered internal corrosion resulting in to frequent interruptions due to heat exchanger failures.

          4.7.4  Policy Environment

While there were many and great advantages in administered price system to provide cheap fertiliser to the farmers and compensate the manufacturer with reasonable cost of production, the system did not provide incentive to the manufacturer to upgrade the technology. Capital expenditures for up-gradation were difficult to get reimbursed and any efficiency gains after up-gradation were moped up under pricing mechanism.

4.7.5  Management Practices

Awareness towards the energy conservation was low during the decade of 1980’s.  Management emphasized on increasing production by improving on-stream factor.

The energy consumption levels on all India level are much improved now due to better operation & maintenance practices and innovation and modernization of old plants.  The energy savings already achieved by the industry at the current production level is equivalent about a million tonnes of fuel oil for a year (for the fertiliser industry as a whole) when compared with 2002-2003 energy consumption (for the current production) and 87-88 levels of energy consumptions. Presently, the Indian gas based plants compare well with the American gas based plants.