Wednesday, 15 February 2017

PROCESS TECHNOLOGY Use monitoring programs to improve ammonia plant uptime

Use monitoring programs to improve ammonia plant uptime
Trending processing information – pressure, temperature – provides ‘clues’ on the reliability of critical equipment and catalysts
M. P. Sukumaran Nair, Travancore Cochin Chemicals, Ltd., Cochin, India

Plant profitability is directly related to process and equipment reliability and on stream time. Many dynamic forces constantly impact the processing environment and affect equipment conditioning. What are the options that engineers can use to improve plant operations? Online monitoring is one method to measure plant condition and identify processing and equipment problems before catastrophic failure occurs.
In the following case history, a suggested monitoring program is outlined for an ammonia plant. The checklist provides a quick surveillance guide for critical systems. Monitoring programs are very important, but these initiatives must also incorporate process design, engineering and operating philosophy and practice.

Dynamic nature. No plant can be regarded as a monolith cast out of steel or stone. It is a conglomerate of different systems composed of varying thermodynamic nature, to achieve a definite objective such as manufacturing designated products at specified quality and quantities in a cost-effective manner. Thus, monitoring or tracking programs are essential to maintain processing units within the desired limits for optimum efficiency and productivity. The ammonia plant is no exception to this rule. A well-defined performance-monitoring program is an essential support system for smooth operation.
Performance monitoring is a technical service function – an activity provided by the process-engineering department. Currently, the specific energy consumption in a modern natural-gas reforming ammonia plant using state-of-the-art technology is only 20% above the theoretical minimum. New developments in ammonia technology are aimed at reducing investment costs and raising operational reliability. In operations, reliability is a function of identifying problem areas, evaluating performance trends, strictly complying with all safety and environmental requirements and achieving high on stream efficiency. The objective for monitoring programs is to ensure reliability in operation through periodic performance evaluation, identifying and offering solutions to problem areas and reducing inefficiency and losses.

Data collection. Collecting plant-operating data is the first step. This may be done by using a specifically designed questionnaire. It should address essential information as relating to daily plant operations. The concerned process engineer should visit the plant daily with the above format and get the required information. Plant test runs may be conducted on a quarterly basis for data collection. This information may be used to evaluate the varying process parameters. During such test runs, the plant loading should be kept at steadystate conditions, with operating data collected round the clock. The production department should receive sufficient notice before doing unit test runs. Operations will need time to have instruments calibrated, arrange for special laboratory analysis, etc.
From the collected data, average hourly feedstock and fuel consumption, steam import / export, cooling water and boiler feed water makeup, power consumption, instrument air consumption, ammonia and carbon dioxide production, etc., can be estimated. Detailed analysis of liquid effluents and stack emissions should be measure as part of the monitoring program. New data can be compared with past results, as well as, design parameters from process flow sheets. Such preparatory work can eliminate gross errors found in the detailed evaluation phase. If abnormal indications are noted, then the readings should be rechecked before assessing the real situation. Also, if an analysis is done to see the effect from other parameters either directly or indirectly associated with a particular observation, then the logic of its variation may be understood. Table 1 lists the performance indicators for a typical ammonia plant. We will discuss the individual indicators.
Performance indicators for a typical ammonia facility
Table 1. Performance indicators for a typical ammonia facility
Table 1. Performance indicators for a typical ammonia facility
Plant loading and material balance
Furnace and fired-heaters efficiency
Catalyst performance
Heat exchanger fouling
Reformer efficiency
Compressor efficiency (polytropic)
CO2-removal efficiency and heat balance
Boiler efficiency
Specific consumption figures
Cooling tower efficiency
Steam and condensate balance
Gas loss from the plant
Overall plant pressure drop
Insulation survey
Effluent analysis

Fired heaters. Feedstock pre heaters upstream of desulfurization reactors and process air preheater are natural-draft fired heaters with liquid / gas burners at the bottom. Efficiency of these heaters is about 85%; hence, tremendous quantities of heat are wasted through the stack with flue gas. The arch draft is usually maintained around 2.5 – 3 mm WG so that air leakage is minimal. Heat from the flue gas can be recovered by installing a secondary feed preheater coil in the convection zone above the existing ones so that the flue gas is cooled to 150°C. Proper combustion inside the fired heater is achieved by adjusting the primary / secondary air to the burner through the respective air registers, plugging all air-entry sources into the heater and maintaining a good flame geometry. The burners should be tuned to get a stack analysis with oxygen content of not more than 4% – 5% in the flue gases in the case of oil firing and 2% – 3% for gas firing.

Primary reformer. The endothermic steam reforming reaction proceeds with an increase in volume and according to LeChatlier’s principle, a higher pressure will shift the equilibrium to the left. This is compensated and the same residual methane is obtained at the outlet by raising the temperature. Two limiting factors of reformer tube design are availability of materials capable of withstanding mechanical stress from internal pressure and thermal stress due to high temperature, and heat flux.
The primary reformer is laid out in two sections – a radiant firebox containing burners and tubes packed with reforming catalyst, and a convection section containing a series of heat exchangers that form the flue gas heat recovery train. Temperature inside the firebox is maintained so that the tube wall is around 730°C – 930°C and about 50% of the heat is absorbed by the entering process feedstock at 455°C – 565°C to facilitate the endothermic reforming reaction.
The performance of the primary reformer is linked to several factors that are critical for the ammonia plant’s efficient operation. They are: heat flux in the radiant section, burner disposition, flame impingement on the tubes causing hot spots and bands, furnace draft, voidage in the catalyst packing, plugged tubes, channeling of flow, carbon deposition, loss of catalyst activity and increase in pressure drop. The heat flux through the tube walls in conventional reformers is about 60,000 W/m2, and in certain cases, it may reach 75,000 W/m2. Thin-walled reformer tubes allow a higher heat flux at a lower wall temperature. Monitoring the tube-skin temperature of reformer tubes is very important and it is useful when diagnosing plant problems before tube failures occur. Optical or infrared pyrometer can be used to measure tube-skin temperatures. Often, the tube-wall temperature trend is more important rather than absolute temperatures. For data collection purposes, a simple optical pyrometer can be used to collect information on a routine basis. Tube life, as designed, is a maximum of 100,000 hours on the basis of creep rupture data. The service life of the tube is shortened considerably from several factors. Premature tube failure may occur due to catalyst poisoning, loss of steam, thermal shocks, overfiring during startup, accidental water carryover to the hot tubes, frequent startups and shutdowns, etc.
Loss of catalyst activity from sulfur poisoning reduces the rate of reaction and lowers heat absorption from the tubes and thus, may create hot bands. Under normal operating conditions, the tubes are the hottest parts and operate at or near structural stability limit of the tube metallurgy. Hence, these hot bands can have disastrous consequences on tube life. Hot bands on tube outside surface may also result from improper catalyst loading, which creates voids in the catalyst layer and lowers heat absorbed by the gas. To avoid such situations, catalysts should be loaded very carefully so that the pressure drop across individual tubes is almost uniform. Improve reformer operation to avoid direct impingement by flames on the tubes. Options include observing flame pattern, regular cleaning of burner nozzles, etc., especially when using liquid fuels.
Inspect the mechanical condition of the tube supports frequently to ensure free expansion of the tubes occurs without bowing. The overall thermal efficiency of the primary reformer may be as high as 95% in a top-fired furnace. The radiant section efficiency varies from 45% to 49%, based on the lower heating value (LHV) of the fuel and the balance heat recovery that occurs in the convection section. The radiation efficiency of side fired furnace is lower than that of a top-fired system.

Catalyst performance. The primary and secondary reforming reactions are equilibrium controlled; the operating temperatures as indicated by the outlet temperature are constant. Under these conditions, the approach to equilibrium can be used to measure catalyst performance. The approach to equilibrium is the difference between the temperature at the outlet of the catalyst bed (Tout) and the equilibrium temperature corresponding to the gas composition (Teq). When calculating the approach to equilibrium, two factors can affect this calculation – variations in the gas composition during sampling and changes in the outlet gas temperature. An approach to equilibrium of 5°C is considered very good. Also, the trends in the change to equilibrium are more important when assessing catalyst performance. However, difficulties, such as radiation effects, thermocouple failures and even thermocouple location, can adversely affect accurate measurement of the gas outlet temperature.

Pressure drop. Pressure drop across the inlet and outlet of the primary reformer may result from catalyst breakage, disintegration and deformation. Aging and faulty operation may occur from carbon overlay, sulfur or arsenic slip or degradation of support materials and these are contributors to poor performance. Pressure drop should be measured on a regular basis as part of reformer-performance monitoring program using the same standard pressure gauge at both ends or by connecting a DP cell across the reformer. A minimum pressure drop is necessary to ensure even process gas distribution through the catalyst tubes. Any increases over the design limit will reduce plant loading and necessitate higher compression power requirement for synthesis-gas compressor.

Methane slip, also known as the equilibrium concentration of methane at the reformer outlet, is a function of temperature, pressure and steam / carbon ratio for the reforming process. Major reactions occurring in the reformer are endothermic-reforming and water-shift reactions. The former is favored by high temperature and low pressure and latter by low temperature, but is unaffected by pressure changes. To maximize the overall economic considerations, the reformer is usually operated at higher temperatures, as well as, high pressures. The stoichiometric requirement of steam is one mole per carbon atom for the methane reforming reaction. However, in practice, it is observed that the catalyst may promote carbon formation under such conditions. Minimum practical steam / carbon ratio for methane reforming is 1.7, and that for naphtha reforming is 2. However, large reformers use a steam / carbon ratio of 3 – 3.5 for optimum reaction. The equilibrium gas composition can be recalculated from the inlet steam / carbon ratio and operating parameters of the reformer to check the deviations from designer’s intentions.
Feedstock quality. For the feedstock – natural gas, important parameters affecting reformer operation are gas pressure, composition and net-calorific value. Reduced pressure from the supplier will cause fluctuations in gas intake and consequently, reduce plant load. A lower net-calorific value of gas requires consuming more gas to maintain the same production levels. Often, the front must be operated at higher loads to supply the requisite CO2 quantities to the urea plant.
For naphtha reformers, feedstock quality has a direct impact on throughput and catalyst service life, as well as, that of the tubes. Typically, reformers are operated well below the maximum conditions allowed for naphtha loading (weight of naphtha in kg/l of heated catalyst). Because of other considerations such as equilibrium approaches and maximum tube-wall temperature, ammonia-technology licensors may recommend the following maximum loading for naphtha with a final boiling point of 180°C and a maximum of 12% aromatics.
Exit temperature, °C
Maximum loading, Kg/l




The reductions in maximum loadings from higher final boiling point (FBP) and aromatic content are listed in Table 2.

Table 2. The reduced maximum loading from higher FBPs with resulting aromatic content
Final boiling point
(FBP), °C
Reduction of maximum
loading, %





Aromatic content, %




Secondary reformer. Secondary reforming is the partial oxidation of residual methane in the primary reformed gas. Performance of this reactor is monitored by checking the gas outlet temperature, methane slip and approach to equilibrium (methane / steam). Measuring the outlet temperature is more difficult than for the primary reformer due to radiation effects, thermocouple failures or poor gas distribution. Catalyst-bed temperature profiles are difficult to quantify. The outlet temperature is measured in the transfer line to the reformed gas boiler. Heat loss from the bed outlet through the fused alumina packings and gas collector pipe to the thermocouple point lowers the outlet temperature.
If the burner is poorly designed or in poor condition from overheating or damage, air distribution will be uneven. Thermocouples used to measure temperature in this region are made from materials such as Inconel 600; yet, thermocouple failure is common. By analyzing the outlet gas, it is possible to calculate the theoretical reaction temperature, assuming adiabatic conditions. This calculated value, which is invariably higher by assuming no heat loss, may be used to check the measured outlet temperature. The gas outlet temperature and its methane content may provide an accurate estimate of the reactors performance, catalyst activity and gas distribution. Pressure drop across the reactor gives an indication of the condition of the vessel’s refractory lining, fusion of the top layer of catalyst, etc. Excessive steaming from the water-cooled jacket indicates refractory failure. In most reactor failures, it is burner failure that causes damage to refractory linings and catalyst.

HT converter. The thermodynamics of water-gas shift reaction suggest that the equilibrium constant decreases with temperature; and by being reversible and exothermic, the equilibrium CO concentrations at elevated temperature are high. The reaction is virtually unaffected by pressure and raising the steam / gas ratio will improve conversion of CO to CO2. As the reaction proceeds, the heat of reaction increases the operating temperature and thus, restricts further conversion. To achieve lowest CO levels at the outlet, the reaction must be maintained in two or more stages with intermediate heat removal. Important parameters to monitor for this reaction are catalyst-bed temperature profile, approach to equilibrium (water-shift), pressure drop across the bed and CO content of the exit gas.
A new catalyst charge is usually operated at the lowest temperature to preserve high activity as long as possible. As the catalyst’s activity decreases, the inlet temperature is raised. At end-of-run (EOR) conditions, the catalyst is operated at temperatures limited by the vessel design – to maintain minimum desired CO slip. Although the increase in inlet temperature with decrease in catalyst activity is gradual, the catalyst-bed temperature profile should not show marked variations, since the deterioration of catalyst activity is gradual throughout the entire bed. The approach to equilibrium is easy to calculate provided that all process parameters are held constant during data collection. This is critical, because in actual practice, the trend is to optimize operations to achieve a lower CO leakage by increasing the catalyst-bed’s operating temperature. Under this situation, the catalyst service life is directly linked to higher inlet temperatures and the system restrictions prevent further temperature increases to keep CO leakage at a desired level until catalyst change out is necessary. The reactor may experience a higher pressure drop due to catalyst fragmentation during use, carry over of dust and steam-volatile compounds from the reformer side, condensate / steam leaks from heat exchanger or boiler tube failures, etc.
Condensate impingement may also break catalyst pellets into fines. Currently, a nitrogen blanket is placed over the top layer of the catalyst to ease pressure drop across this bed.

LT converter. The operating efficiency of modern ammonia plants heavily depends on the performance of low temperature (LT) shift catalyst. Located at the downstream of high temperature (HT) shift, this is the last reaction stage to convert CO into CO2 and hydrogen. Reducing the CO content from LT shift reactor by 0.1% directly contributes to a gain of 1% to 1.16% of the rated ammonia production. Due to the reversible and exothermic nature of CO conversion and maintaining the lowest CO content in exit gas, the latter stage of the reaction is carried out at low temperature (200°C – 250°C).
A highly active copper catalyst is also used to keep the equilibrium CO concentration to less than 0.3% at the outlet. Using this practice, the complicated process steps such as copper-liquor wash and liquid-nitrogen wash are eliminated; also, it makes the subsequent methanation stage economically viable. It is the catalyst’s performance rather than reactor design that significantly determines the efficiency of LT shift reaction. The LT shift catalyst is more expensive in terms of catalyst consumption per ton of ammonia produced.
Essential requirements for efficient operation are stable and high catalyst activity, resistance to poisoning and mechanical strength. Loss of catalytic activity is known as poisoning, and loss of active surface area is due to thermal sintering. Copper compounds have relatively lower melting points compared to similar components of iron or nickel and, thus, it is necessary to operate nickel catalysts at lower temperatures. To improve thermal stability, alumina is also added to the oxides of copper and zinc during the formulation. Copper has good affinity to sulfur; thus, the poisoning effects are very high. However, in presence of zinc oxide, the bulk of the sulfur is retained by the zinc; the copper surfaces are unaffected and continue to operate satisfactorily.
Due to higher chemical affinity towards chlorine and high volatility of cuprous chloride (melting point – 30°C), chloride acts as a powerful poison for the LT shift catalyst thus, decreasing thermal suitability. The important parameters used for predicting LT catalyst performance are outlet gas analysis, temperature profile and pressure drop across the reactor. Chromatographic analysis of outlet gas usually measures the CO content. This can be counter checked by knowing the temperature rise across the methanator if the CO2 slip from the CO2-removal unit is known. A new catalyst charge could be operated at a lower inlet temperature, i.e., 20°C higher than the dew point of inlet gas (200°C), and achieve an acceptable close approach to equilibrium CO content.
Progressively, as the catalyst deactivates, the same approach is maintained by increasing the outlet temperature. Subsequently, the temperature profile moves down through the bed at a constant rate with the gradual decline of catalytic activity. As the reaction zone approaches the bottom of the bed, the CO slip gradually increases and a catalyst change out is necessary. Monitoring the shift in the hot spot from the temperature profile can enable making accurate forecasts of catalyst life. Any increase in outlet CO content above 0.3% will definitely affect lost production. Although the deactivation of LT shift catalyst is gradual and uniform, it can also be aggravated by upstream plant problems, changes in feedstock quality or inefficiency of feedstock purification catalyst. Water condensation on LT shift catalyst is also dangerous; poison accumulated on the top layers may be washed down to lower parts of the catalyst bed rendering it inactive. It can also cause catalyst fragmentation; thereby, increasing pressure drop across the bed.

Methanation. Methanation of residual carbon oxides is the last stage when preparing synthesis gas for ammonia manufacture. With the development of LT shift catalysts that can substantially decrease CO slip, it is now cost-effective to integrate the methanation stage and replace the troublesome copper-liquor wash and liquid-nitrogen wash stages for synthesis gas purification. Two important monitoring components in methanator operation are the carbon oxide concentrations in the exit gas and the reactor temperature profile. Nickel catalysts that are used have high surface area, stable alumina support and a high degree of nickel dispersion; these factors induce much stability and activity. The methanation reaction is strongly exothermal, giving a temperature rise of 75°C per percent of CO and 60°C per percent of CO2.
The acceptable level of total carbon oxides in syn gas is less than 10 ppm by volume (ppmv). Temperature rise measured across the bed alone cannot be used as performance indicator for methanator. The temperature is increased to reduce carbon oxides from 30 ppmv – an unacceptable level – to <10 ppmv – an acceptable level – are more or less the same though the reverse is true. In combination with the outlet analysis for (CO + CO2), the reactor temperature profile can provide the operator with sufficient information to estimate the catalyst’s performance and remaining service life.
Catalyst deactivation is mainly due to poisoning. The poisoned zone, which is indicated by the shift in the hot spot in the bed, will progressively move down through the bed overtime. As the remaining active catalyst become insufficient to meet the <10 ppmv level of carbon oxides, a catalyst change out may be necessary. The catalyst is resistant to thermal sintering; no appreciable loss of activity is evident when it is occasionally over heated, even to 400°C above the normal operating temperature. In such cases, the reactor design temperature limits the maximum temperature. Sulfur poisoning of the catalyst is a rare phenomenon, as most of the H2S will also be removed in the CO2-removal section. Carryover of arsenic trioxide from CO2-removal units can result in permanent activity loss. Small amounts of moisture in the inlet gas to methanator usually does not cause any concern. However, higher moisture levels – 1% to 2% of water – in the feed can inhibit the reaction.

Ammonia synthesis. The ammonia synthesis reaction between elemental hydrogen and nitrogen is highly exothermic, and equilibrium conversion of ammonia increases with rising pressure and decreasing temperature. Selecting the ammonia-synthesis loop pressure is often a trade-off between the advantages of higher equilibrium concentration of ammonia at very high pressure and the rising costs of synthesis-gas compressor and additional capital investments for plant equipment. Plants built in the ’70s and ’80s operate at a loop pressures ranging from 200 – 250 bar. The latest technologies have optimized the loop pressure down to 80 – 150 bar. Although thermodynamics favor a lower temperature for the synthesis reaction, it may be necessary to operate the converter at higher temperatures due to kinetic considerations.
Catalyst activity is very important in this perspective. The best catalyst is the one that will give the higher rate of conversion at the lowest temperature. As the conversion proceeds, the heat of reaction raises the temperature of catalyst bed and the specific reaction rate is increased. At high temperatures, the equilibrium becomes more unfavorable; the dissociation reaction increases and the overall reaction becomes equilibrium controlled. This aspect warrants efficient temperature control and heat removal to maintain the system between units that are set by thermodynamic equilibrium and the kinetics of catalyzed reactions. Thus, in most commercial converters, gas entry and exist of the catalyst beds are maintained around 400°C and 500°C, respectively.

Gas compression. Multi-barrel compressors, driven by HP- steam turbine, are used for synthesis-gas compression along with recirculation of unconverted gas stream from the synthesis loop. Regular monitoring of operating parameters include vibration levels of compressor barrels and turbines, steam consumption, and lube oil makeup along with occasional assessment of polytropic efficiency. Close vigilance of these conditions should ensure smooth operation of the machine.
Drive-steam purity is an important factor. Sometimes, colloidal silica is carried along with high-pressure steam and deposited on the nozzles, thereby, restricting steam input with consequent increases in turbine back pressure and reduced plant load. Performance of interstage coolers and efficiency of moisture removal separators also affect the performance of the synthesis-gas compressor. Frequently check the antisurge valve to ensure that it is not passing and, thus, resulting in power loss that is absorbed by the machine. Another important factor is controlling the gas circulation rate. Variations in the circulation rate can alter the composition of the reaction mixture. Higher ammonia content in the circulating gas stream can shift the reaction away from equilibrium, and a lower ammonia concentration in the circulating gas will favor the reaction.

Synthesis loop. The H2/N2 ratio at the inlet of the converter has a profound effect on plant operability. For a constant rate of makeup gas intake and purge, the operating pressure required to achieve the conversion as a minimum when the H2/N2 ratio is three. Any deviation from this ratio increases the average pressure in the loop and the rate at which this increase takes place also raises with the deviation of H2/N2 ratio from three. A small change in H2/N2 ratio in the makeup gas results in a much greater change in the H2/N2 ratio in the reactor and other streams of the loop. Thus, ratio control of the makeup synthesis gas, through air addition in the secondary reformer or during gas purification by liquid-nitrogen wash assumes greater importance and warrants close control. Experiments carried out in commercial converters show that the maximum conversion is obtained with a ratio in the range of 2.5:1 to 3:1, depending on operating conditions and limits of overall optimization.

Influence of inerts in the loop. Synthesis gas from steam reforming plants generally contains about 1% methane and 0.3% argon. Conversely, partial oxidation plants using liquid-nitrogen wash for final gas purification yield a very pure synthesis gas containing not more than 0.01% methane and argon. These constituents are non-reactive, build up in alarming proportions in the circulating gas and enormously lower the partial pressure of the reacting gases.
Controlling inert concentration in the loop circulating gas is important and is done by maintaining a constant purge. The higher the purge rate, the lower the inert concentration in the loop. This is favorable for the reaction, but hydrogen loss from the loop will be greater leading to lost production. The inert level is controlled at 12% – 13% in plants where the synthesis gas comes from the steam-reforming route.
Ammonia and hydrogen from the purge gas is usually recovered and the latter is recycled back to the loop. Hydrogen recovery may be done cryogenically or by using permeable membranes or molecular sieves. Purge-gas hydrogen recovery raises ammonia production by 4% – 5%. To maintain the H2/N2 ratio without installing purge-gas hydrogen recovery unit, it will be necessary to add more air in the secondary reformer than is required for a 3:1 stoichiometry. In such cases, the fresh makeup gas may have a H2/N2 ratio of 2.8:1. For a given makeup gas composition, the recycle and purge ratio and ammonia production remains constant. With higher inert levels, loop pressure may build up; consequently, it may be necessary to raise purge rate and loop circulator to maintain the desired production level.

Converter inlet ammonia content. Complete removal of ammonia contained in the converter outlet gases is not possible in subsequent condensation and separation stages. Normally, 1.5% – 5% ammonia will remain in the recycled gas, which will adversely affect the equilibrium behavior of the converter. Factors that contribute to elevated ammonia concentrations in the gas inlet of the converter are the ammonia-content exiting the water-cooled condenser, conversion per pass, and the vapor / liquid equilibrium between ammonia and synthesis gas at the operating conditions. These factors can be the direct outcome due to fouling of water condensers, increased ammonia content in the converter outlet, separator failures, higher separator temperatures, a leak in the internal exchangers of the catalyst basket, gas bypassing, etc. These symptoms are evident from deviating operating conditions and can be confirmed by chemical analysis of various streams.

Catalyst poisoning. Ammonia synthesis catalysts are very sensitive to poisoning that can be either temporary or permanent. Temporary poisoning occurs from exposure to oxygenated compounds such as water vapor, carbon monoxide or carbon dioxide, etc. These compounds may be carried along with the circulating gas. Activity loss due to these compounds is largely reversible. If the exposure does not exceed a few days, the activity can be resumed by simply using poison- free synthesis gas at slightly elevated temperature for several days. If the separators malfunction and liquid ammonia is carryover to the catalyst, then the catalyst may be poisoned by water and CO2 dissolved in the liquid ammonia.
Permanent catalyst poisoning occurs from the presence of sulfur, phosphorus, chlorine or arsenic. This exposure causes irreversible activity loss. Sulfur poisoning may occur via lube-oil carryover from reciprocating compressors. To reduce this risk, use lube oil with a low sulfur level – 0.1% to 0.15%.
In plants where arsenic trioxide is used as a promoter in the CO2-removal stage, there is a potential for inadvertent introduction of arsenic into the gas stream. Chlorine compounds may also inhibit catalyst activity by converting it into volatile metallic chlorides. Thermal decomposition of hydrocarbons contained in lube oil and consequent carbon overlay onto the catalyst is another way to foul catalyst active centers. Modern plants use well maintained centrifugal compressors and circulators; hence, problems related to sulfur poisoning are practically eliminated.
Apart from catalysts and catalytic reactors, the performance of the ammonia plant also depend on a number of systems and equipment, some of which are discussed below.

Table 3. Effects from catalyst “poisons”

LTS catalyst poison

Effect on catalyst

Hydrocarbon feed
Steam quench
Lube oil HTS catalyst

Covers the active
copper surface

Hydrocarbon feed
Steam feed
Lube oil
Air to secondary reformer
Enhances growth of
copper crystals,
besides covering
the active surfaces

Steam quench
Upstream catalysts
Brick lining
Bed support materials
Physically blocks

CO2 removal
Normally not

Boiler feed water
Normally not

HT converter
Normally not

Heat exchangers. Fouling of heat exchangers is a severe problem especially in sections such as naphtha hydrofiners, CO2-removal, compressor intercoolers, synthesis-gas cooler, water condenser, etc. Proper cooling water treatment and using side-stream filters will inhibit fouling. For critical exchangers, the overall heat transfer coefficients may be calculated from operating parameters at regular intervals and compared with the design values. Such evaluations will assess and identify the deterioration over time. A periodic cleaning or back-washing program for these exchangers may be necessary.

Compressors. Polytropic efficiency of centrifugal compressors is the best performance indicator. The deterioration trend of the compressor’s operating efficiency can be predicted from a plot of polytropic efficiency over time.

CO2 removal unit. Loss of CO2-absorbent solution, excessive LP stream consumption, corrosion of equipment, etc., are problems usually encountered in CO2-removal units. To ensure trouble-free operations, monitor pressure drop across the absorber to avoid solution carryover into the gas stream, solution analysis, thickness measurement of critical pipelines, etc.

Pressure drop. The overall plant pressure drop is a general indicator of plant performance. It will identify the sudden collapse of linings, refractories, tube leaks, malfunctioning of control valves, choking of demisters, etc. A trend analysis of the pressure drop will help operators to identify performance deterioration of individual equipment, especially in areas such as the primary reformer, RG boiler and other catalyst-packed reactors.

Gas loss. Loss of process gas from passing of drain and vent valves, valve leakage, etc., creates a hazardous environment. This situation may be overcome by proper vigilance.

Steam-condensate balance. The steam-condensate balance is an important checklist when tracking steam and water management. Abnormal variations in steam production and consumption are indicative of unsteadystate conditions or equipment failures.

Insulation survey. Upkeep and maintenance of hot and cold insulation is a serious concern both for efficient plant operation, as well as, for energy conservation. Regular inspections of refractory-lined pipelines and vessels, superheated steamline insulation, etc., are essential.

Effluents. Effluent and emission analyses provide a discrete picture of the steady state operating conditions. Abnormal variations in quantities and qualities of effluents and emissions signal deviations from design specifications. The above is a general listing of things that matter operators of ammonia plants worldwide.


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Appl, M., "Modern ammonia technology: where have we got to, where are we going?," Nitrogen, September – October 1992, November – December 1992, March – April 1993.
Buvidas, L. J., "Cut energy cost in ammonia plants," Hydrocarbon Processing, May 1979.
Garg, A., "Optimize fired-heater operations to save money," Hydrocarbon Processing, June 1997.
Houken. J., "Optimizing catalyst performance and change-out schedule," IFDC Seminar, Houston, 1989.
ICI, Technical bulletin – 7, "Monitoring catalyst performance," ICI, Billingham, England.
Low, G., "Revamping ammonia plants – improve process operations and economics," Supplement Nitrogen, January – February 1984.
Madhavan, S., and B. Landry, "Technical audit of existing ammonia plants," AIChE Ammonia Symposium, Minneapolis, 1987.
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