Back-to-Basics: Maximizing Plant Performance through a Better … · 2018. 8. 22. · operation can...

Back-to-Basics: Maximizing Plant Performance through a

Better Understanding of the Basics in Catalytic Reactor

Operation and Problem Avoidance

Catalytic technologies are being further developed and enhanced for the ammonia production process

to enable improved performance and better utilization of our natural resources. This goes in line with

the need to minimize the costs of ammonia production in a competitive market, determined by volatile

or increasing prices of the commonly used natural gas feedstock. Superior performing catalysts are a

very important key to achieving this success. Despite the development of high performing and robust

materials, it is still necessary to have a thorough understanding of the chemical and physical basic

principles of the catalytic reactions involved. This allows the user to achieve maximum performance

and to avoid consequences such as pre-mature change-out or poor catalyst efficiency caused by such

circumstances as poisoning, improper flow distribution, or improper operation. This paper describes

fundamentals of catalytic reactor operation and how their application can help to ensure that

superior catalyst activity can be achieved in various catalytic services of an ammonia plant.

Michelle Anderson

Süd-Chemie Inc.

Scott Osborne

Süd-Chemie Inc.

Introduction

n the ammonia industry, high turnover, flatter and leaner organizational structures, new plant start-ups and restarts have

contributed to changes in the make-up and demographics of the operations and technical personnel staffing. These changes in the experience and skill-set base have resulted in a loss of long-term knowledge and lead to a tendency for history to repeat itself in the form of incidents of equipment and catalyst damage during operation and upset conditions.

Despite advances in catalyst technology that have enhanced ammonia plant production and energy utilization, a thorough understanding of the chemical and physical principles of catalytic reactions is required to maximize performance and avoid consequences such as premature change-out, poor catalyst efficiency and even equipment failure. This paper focuses on three particular areas where industry experience has shown to have a significant impact on a plant’s performance. These areas of discussion include; (1) chemistry, causes, consequences and avoidance of carbon formation in the steam reformer, (2) the importance of proper flow distribution in various types of catalytic reactors, and finally (3) some examples of the

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significant negative impact of improper plant operation on catalyst and equipment.

Carbon Formation in Primary

Reformers

Steam reforming is required to convert hydrocarbons into hydrogen which is an important step in the process used for the manufacture of ammonia. The primary reformer is the main piece of equipment required to accomplish this goal and represents the largest expenditure in terms of capital and on-going energy costs. Optimum performance of the reformer and installed reforming catalyst is critical to ensuring high plant productivity and efficiency. Poisoning, fouling or mal-operation can adversely affect the catalyst’s performance which can lead to costly equipment failure.

Primary Reformer Reactions

The main reactions in the primary reformer are the following: Steam reforming of hydrocarbons CnH2n+2 + n H2O ↔ n CO + (2n+1) H2 (1) ∆Hr > 0 Steam reforming of methane CH4 + H2O ↔ CO + 3 H2 (2) ∆Hr = 88700 Btu/lb-mole (206 kJ/mol) Water-gas shift CO + H2O ↔ CO2 + H2 (3) ∆Hr = -17,600 Btu/lb-mole (-41 kJ/mol) Overall, the process is endothermic. A lower CH4 leakage is favored by higher exit temperature, higher steam-to-carbon ratio, and lower exit pressure. The standard primary reformer system is comprised of a tubular furnace where the feed

stream passes over a catalyst packed in multiple banks of externally heated tubes. With improvements in tube metallurgy, primary reformers operate at pressures up to 870 psig (6,000 kPa) and tube wall temperatures up to 1850oF (1010oC). The average heat flux may be as high as 30,000-35,000 Btu/hr-ft2 (94,000-110,000 W/m2). Firing is usually controlled such that tube wall temperatures are maintained at values that will give a reasonable tube life. By design and industry practice, maximum allowable tube wall temperatures are set to give an in-service life of 100,000 hours.

Carbon Formation Reactions

With no steam present and at normal reformer operating temperatures, all hydrocarbons will decompose into carbon and hydrogen via the following reaction: CH4 ↔ C + 2 H2 Thermal Cracking (4) Cracking reactions are thermodynamically favored at high temperature. In the presence of steam, also gasification reactions occur over the primary reformer catalyst as; C + H2O ↔ CO + H2 (5) Operating the reformer at conditions that drive this reaction can prevent the accumulation of carbon deposits. Should conditions exist which lead to hydrocarbon cracking, the heavier hydrocarbons in the feed will crack first. As this occurs, the active sites of the catalyst become masked, resulting in less reforming reaction and higher gas temperatures which further increases the tendency for cracking and coke deposition. Common causes of carbon formation and deposition are low steam/gas ratio operation, slugs of heavy hydrocarbons in the feed, sulfur poisoning, and temperature excursions resulting from poor firing control. Higher catalyst

260 2011AMMONIA TECHNICAL MANUAL


activity results in lower gas temperature during the period that methane concentration is still high enough for cracking, thereby reducing the carbon formation potential. The overall effect of carbon formation on the primary reformer is reduced conversion, increased pressure drop and increased tube wall temperatures which, if severe, can be seen as ‘hot bands’ on the tubes. Not only does the carbon deposit block active sites, its formation within the pores of the catalyst can cause catalyst breakage.

Monitoring Program

During normal operation of a primary reformer, a routine monitoring program assists in optimizing the catalyst’s performance as well as providing early detection of any issues that can lead to catalyst or tube failure. Three performance parameters are typically trended during operation; calculated CH4 Approach to Equilibrium (ATE), tube wall temperatures (TWT) and relative pressure drop.

Approach to Equilibrium

Reformer catalyst activity is commonly expressed as CH4 ATE, which is the difference between the measured catalyst outlet temperature and the calculated equilibrium temperature for the observed methane leakage. An increased ATE or methane slip from the front-end results in a higher purge rate from the synthesis loop and increased energy consumption for the plant. While operating the primary reformer at higher exit temperature can mitigate this effect, the cost of increased fuel requirements and the impact of hotter tube wall temperatures must be considered when choosing to operate in that non-optimal mode.

Tube Wall Temperatures

Tube costs are a significant item in the overall economics of a reforming plant. The goal is to operate the reformer under conditions that result

in the lowest possible tube wall temperatures (TWTs) consistent with satisfactory reformed gas quality. Even a slight increase in the tube wall temperatures will have a drastic impact on the tube life. A TWT increase of as little as 18°F (10oC) may result in up to a 30% shortened lifetime of the reformer tubes. The use of a high activity / shape optimized reforming catalyst results in lower tube wall temperatures and a significant cost and efficiency advantage to a producer. As an example, a 1200 STPD (1088 MTPD) plant in North America was able to realize such an advantage by installing Süd-Chemie’s high activity ReforMax® reforming catalyst. The Kellogg designed primary reformer was initially designed to produce 600 STPD (544 MTPD) of ammonia. As a result of several expansions, the reformer now operates at a high heat flux of around 35,000 Btu/hr-ft2 (110,000 W/m2). Figure 1 shows that the mean TWT is 100°F (55oC) cooler with Süd-Chemie’s catalyst than with the previous charge.

Figure 1. Comparative average TWT surveys of

two charges of catalyst at a 1200 STPD

ammonia plant

Further, the maximum tube wall temperature was also lowered and the number of tubes with TWTs exceeding the temperature limit of 1700 °F (927oC) went from 5 to zero which significantly reduced the risk of potential tube failure. Figure 2 shows the number of tubes operating above 1700 °F (927oC) during the life

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of the previous charge compared to no such tubes for the Süd-Chemie catalyst charge.

Figure 2. The number of instances where the

maximum TWT exceeded 1700 °F during the life

of the competitor catalyst charge

Given a particular catalyst charge, effective monitoring of the TWTs and balancing of the firing of the reformer furnace can extend the life of the tubes and optimize the performance of the catalyst. TWT measurement studies are typically done with an infra-red optical pyrometer or via infra-red thermography corrected for emissivity and reflected radiation. The measurements are usually taken in the region of highest heat flux as this often coincides to the region with the highest TWTs. A survey measures and plots the TWTs to aid in locating areas that require burner adjustments. Furthermore, TWTs exceeding tube design limits can be identified for further investigation and closer monitoring.

Figure 3. IR thermographic image of a row of

reformer tubes

Figure 4. Plotted TWT surveys showing the heat

distribution across the reformer

Pressure Drop

The impact of increasing primary reformer pressure drop is reduced plant efficiency caused by the lowering of the suction pressure at the synthesis gas compressor and the resulting higher horsepower requirements to maintain the desired loop pressure. Depending on pressure relief valve set-points, a higher reformer ∆P may also require lowering of plant rates if operating near those limits. Through Süd-Chemie’s proprietary CATTRENDS program, calculated ATE and relative pressure drop are trended over time. Carbon deposition on the catalyst will manifest itself as increasing ATE, increasing relative pressure drop and increasing TWTs. In severe

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cases, visual hot bands on the tubes will also become evident.

Operational Control Factors for Minimizing

Carbon Formation

To prevent carbon formation, a minimum steam to carbon (S/C) ratio must be maintained at all times. This value is contingent on the composition of the feedgas, firing rates, operating pressure and incoming feed temperature. Most carbon formation incidents occur, however, during transient conditions such as start-up, trips or shutdowns. Modern operating plants typically have a low steam to carbon ratio trip interlock that cuts out the feed-gas to protect the reformer during upset conditions. Additionally, a properly installed minimum stop on the steam valve ensures that steam continues to flow through the reformer in an emergency trip to sweep all feed-gas out of the reformer. Complete isolation of the feed-gas valve during a trip or shut-down is necessary before steam cut-out. Controlled start-up and shutdown procedures and emergency interlocks are designed to avoid carbon formation. The objectives are to: 1. Prevent over-firing and thus temperature

excursions during rate changes. This typically considers the following operating parameters; fuel flow, fuel pressure, flue-gas temperature, process gas pressure, steam flow, feed-gas flow, combustion air flow, exit process gas temperature and the number and location of burners lit. Many times a manual start-up rate guide is used and in some cases an algorithm with interlocks guides the operator during the start-up.

2. Maintain a high S/C ratio during rate

changes. This is typically achieved by

increasing the steam flow before the feed-gas on start-up and reducing the feed-gas before the steam on shutdowns.

3. Ensure sufficient and even flow distribution

throughout the entire reformer. Before feed-gas is introduced to the reformer during start-up, the steam flow should be sufficient to ensure that all tubes have similar operating conditions i.e. even flow and even heat through all tubes.

4. Ensure sufficient steam flow exists for

adequate mixing of steam and feed-gas well before the tube inlet manifold.

Visual monitoring of the tubes and burners is imperative in detecting any issues during the start-up process.

Sulfur poisoning as root cause for carbon

formation

Sulfur has an adverse impact on reforming catalyst activity and as a result is removed from the feed-gas to levels less than 0.1 ppmv via an upstream desulfurization step. The sulfur reacts with the reforming catalysts active nickel component according to Ni + H2S ↔ NiS + H2 (6) to form an equilibrium condition over the active sites of the catalyst. Sulfur is not a permanent poison and when removed from the feed, the catalyst should in principle regain the original activity. However, sulfur poisoning can deactivate the catalyst to the extent that less reforming is done and gas temperatures elevate causing high tube wall temperatures. Even small amounts of sulfur in the feed can significantly influence the tube wall temperatures as shown in Figure 5.



Figure 5. The impact of Sulfur concentration in

feedgas on the tube wall temperature in the

reformer If the temperature rises to the point where carbon formation occurs, pressure drop will increase and the catalysts performance will be further compromised. In a 600 STPD (544 MTPD) design ammonia plant, the desulphurization unit was bypassed due to vessel design concerns. During this time, Sud-Chemie’s high strength ReforMax® primary reformer catalyst experienced inlet sulfur levels of up to 3 ppmv for nine months. As expected, catalyst poisoning occurred which resulted in some reversible activity loss in the 30-40% top inlet tube location. The effects of the poisoning can be seen in Figure 6.

Figure 6. The hot-bands seen on reformer tubes

after exposure to high levels of sulfur in the

feedgas

The plant was subsequently brought down and a successful steam-air regeneration procedure was done with the assistance of Sud-Chemie’s technical support to remove (i) the carbon that had formed on the catalyst and (ii) the sulfur that had caused the carbon formation. On re-start, the reformer experienced essentially full recovery to pre-incident performance levels and tube wall temperatures as seen in Figure 7.

Figure 7. The reformer tube appearance

following a decoking procedure

The ReforMax® catalyst withstood the arduous conditions during this high sulfur exposure time-frame and continued to provide the full expected design life-time. Given the impact of sulfur poisoning on the performance of the reformer, the choice of high activity desulfurization catalyst should not be under-estimated. Proper monitoring of the total sulfur slip from the pre-treatment section ensures that the reforming catalyst is protected.

Flow Distribution in Reactors

The design of any catalytic reactor needs to consider effective flow and distribution of the feed-gas through the catalyst bed volume. The catalyst performance is predicated on this assumption and any mal-distribution of gas flow can lead to bypassing/channeling within the bed and ineffective utilization of the loaded catalyst volume. This manifests itself in higher

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calculated ATE in primary and secondary reformers, higher CO leakage in HTS and LTS reactors, higher COx from methanators and reduced conversion from ammonia converters. All of these non-optimal conversion impacts lead to less efficient and more costly operation of the NH3 plant. Using the results of a study carried out to examine the impact of such operation, the cost impacts can be converted to real dollar values to demonstrate the large impact they can have on overall plant profitability. This study was based on a 1500 STPD (1361 MTPD) ammonia plant with membrane purge gas recovery [1]. Assuming a production rate of 1500 STPD, US$2.70 per MMBtu natural gas and U.S. $450/ton ammonia, the table below shows the results in terms of potential costs.

Change in Performance

Change in Energy Cost

(USD per annum)

Lost Production (USD per annum)

Primary Reformer

10 F increase in ATE 15,000 170,000

Secondary Reformer

10 F increase in ATE 30,000 370,000

HTS 1% increase

in CO leakage

56,000 590,000

LTS 0.1%

increase in CO leakage

160,000 2,000,000

Table 1. Impact of changes in performance of

selected front-end catalysts on energy costs and

lost production in a 1500 STPD ammonia plant

with membrane purge gas recovery

As shown in Table 1 the increase in costs associated with lost production and – if the lower production can be compensated by a higher energy input – in energy costs can be quite significant for these different scenarios. It should therefore be apparent that proper initial design, procedures and operation is imperative to ensure stable and optimal catalyst performance from SOR to EOR. There are

various potential causes of flow mal-distribution in a catalyst bed and these include gas distributor issues, vessel design issues, movement of catalyst during operational upsets and initial loading techniques. The flow distribution in reactor design, especially in axial/radial synthesis converters, is critical to achieving expected performance. As an example, in an 800 MTPD (882 STPD) ammonia plant, pressure drop issues in the ammonia synthesis converter were causing production losses of 100 MTPD. Ammonia Casale successfully re-designed the baskets in this converter with the aid of sophisticated fluid dynamic simulations. Using this tool and their converter expertise, they were able to specifically develop distribution devices to ensure that the axial and radial flows were distributed such that the full potential of the catalyst was achieved [2]. The plant successfully started up in 2008 and has had steady operation at currently 900 MTPD (993 STPD). Movement of catalyst at the top of a bed caused by poor feed distributor designs, inadequate hold-down, burner issues in a secondary reformer or plant upsets can create flow mal-distribution or channeling issues and result in changes in performance. Such an event was experienced in a 1750 STPD (1587 MTPD) ammonia plant. During operation, a change in secondary reformer performance was seen during routine monitoring as evidenced by an increase in methane leakage from around 0.3% to over 0.5% (Figure 8). On shutdown it was found that there was mixing of the catalyst and inert balls which were pushed up toward the walls of the vessel (Figure 9). The root cause was traced back to an operational upset which caused high gas velocities and bed movement creating unevenness/flow distribution issues in the bed which subsequently compromised the overall performance.



Figure 8. Secondary reformer exit methane slip

over time

Figure 9. Photo of the surface of the secondary

reformer showing mixing of catalyst and inert

balls and bed disruption

The purpose of any initial loading of catalyst and support media is to achieve a uniform density throughout the bed to ensure proper flow distribution. In fixed beds cross-sectional uniformity is also desired. The primary reformer is one of the most intensive and time-consuming reactors to load in an ammonia plant and carries severe penalties if tubes are loaded improperly. With no physical means for balancing the feedgas flow through tubes, catalyst with non-uniform loaded density results in non-uniform flow rates which in turn cause some tubes to operate hotter than the average tube temperatures. Additionally, catalyst bridging can occur when particles

interlock to form a “bridge” across the tube cross-section. These voids within the catalyst layer result in localized overheating of the tube. Hot tubes fail prematurely and as already established can be very costly. It has been estimated that a 2.5% lower than average tube flow rate can result in an 8-10 °F (4-6oC) hotter tube wall temperature [3]. Further experience has shown that tube wall temperatures in the area of a catalyst void are about 30-50°F (16-28oC) hotter than the surrounding area of the tube. As part of Süd-Chemie’s technical service offerings, we provide support to reformer catalyst loadings. This may include involvement in the preparation of catalyst handler bid documents, review of schedules, loading diagram preparation, monitoring loading vendors, outage and pressure drop verification and advice on matters arising during the process. All of this ensures that the reactor is loaded optimally to achieve optimal performance.

The Impact of Plant Operations on

Catalyst

A number of catalyst failures are a result of poor operation or operational upsets. To illustrate the significant potential negative impacts, a number of different types of examples will be presented for various catalytic services.

Reduced/Pyrophoric catalyst

Several catalyst types are in the reduced metal state during operation and exposure to air can result in spontaneous ignition with extremely high localized temperatures causing catalyst failure, fusing and equipment damage. In an ammonia plant, these catalysts are HDS, HTS, LTS, methanator and ammonia synthesis. Operators are typically aware of this risk during a shutdown and measures are taken to prevent air ingress by isolating and inert blanketing these particular reactor systems. For turnarounds

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and longer shut-down periods, ensuring a continuous positive pressure inert blanket is one routine monitoring item. Discharging of pyrophoric catalyst without first oxidizing it must be done under nitrogen with precautions taken to avoid ingress of air to the vessel by a chimney effect. In cases where repair work is required around vessels without adequate isolation, the risk is particularly high and continuous nitrogen purge is necessary. Caution should also be taken if workers are required to do work in the area where nitrogen is purging. It should be noted that without flow through the catalyst bed, temperature measurements do not accurately represent the bed temperatures nor do they detect localized peak temperatures. They can, however, give some indication or warning of an issue. Complete isolation of the air flow to the secondary reformer is required in a secondary reformer trip to protect downstream catalysts. In one incident at a 1700 STPD (1542 MTPD) ammonia plant in North America, the HTS catalyst performance was greatly affected following an upset during start-up of the secondary reformer. A high pressure air flow condition resulted in air to the HTS for a short duration. The event occurred relatively fast and other events were occurring so it was not even noticed by the operations team. Shortly after stabilizing plant rates, however, the performance of the charge in terms of CO leakage, while acceptable to allow operation, was higher than expected. After approximately a year on-stream, the decision was made to unload the catalyst during a shut-down. Inspection of the reactor showed that the overall outage from the manway had increased by over 3 ft (0.9 m) and that in a particular area of the bed the catalyst particle size was reduced from 6x6 mm to roughly 3x3 mm. The support balls were also found fused to melted screen pieces which in turn were fused to catalyst. For this level of shrinkage and fusion to occur, it is estimated that temperatures must have exceeded 1600 °F (870oC). It was particularly impressive

that despite this significant temperature exposure, the Süd-Chemie catalyst was still able to achieve an acceptable level of performance with no pressure drop issues. (Figures 10 and 11)

Figure 10. Photo of reduced particle size of

HTS catalyst after exposure to extreme

temperatures caused by air introduction to the

HTS reactor

Figure 11. Photo of fused inert ball, floating

screen and catalyst removed from the HTS

following exposure to air and extreme

temperatures

Impact of Water on Catalyst

Water can affect the integrity of all catalysts and in some cases can lead to catastrophic incidents. Water entrainment during start-up of the primary reformer can lead to instantaneous flashing at the hot catalyst releasing enough energy to rupture reformer tubes. Water entrainment can occur when low point drains are plugged or not used during shutdown and start-up.



HTS and LTS are particularly at risk due to their location downstream of the HP boiler, quenches and temperature controls. When water comes in contact with catalyst, catalyst breakage and increased pressure drop results from water vaporizing out of the catalyst pores at a faster rate than the porosity of the catalyst will allow. The HTS catalyst is very susceptible to leaks from the upstream HP Boiler which are almost inevitable during a plant’s life. Süd-Chemie’s HTS catalyst was able to survive major boiler leaks at the 1600 STPD (1451 MTPD) Terra plant [4]. In this case, the HTS was flooded with water from a major boiler failure and a dry out procedure was successfully implemented under the direction of Süd-Chemie. Besides the effect of wetting the catalyst, accumulated boiler solids from boiler leaks can also increase the catalyst’s pressure drop and create poor gas distribution issues. The impact of water and boiler solid carryover can be minimized with a high voidage inert top layer as supplied by Süd-Chemie. During operational upsets, temperature controller swings can lower the inlet temperature of the shift converters below the dew point and wet the catalyst. Operating the LTS at an inlet temperature 30-40°F (17-22oC) above dew point can reduce this potential. Boiler feed water quenches are normally installed upstream of shift converters in order to control inlet temperature as well as lower the CO equilibrium. Atomization is critical to prevent wetting the catalyst and proper design and maintenance of the quench valve / nozzles is imperative. Quench valves also require complete isolation during a shutdown to prevent water buildup in low point piping areas that can be entrained onto the catalyst on restart.

Water Treatment

Much attention is paid to the quality of steam in an ammonia plant for rotating equipment purposes. However steam quality is an important consideration for reforming catalyst. Carry-over and deposition of silica and sodium

on reforming catalyst can act like a permanent poison by physically covering the active sites of the catalyst. This reduces the extent of reforming and leads to poor reformer performance, increased pressure drop and increased TWTs at the top of the tubes. This paper has already discussed the implications of poor catalyst performance and high TWTs. These foulants can also be introduced to the process via boiler feed water quenches upstream of the reformer. A well controlled and monitored water treatment program is necessary for effective catalyst performance.

LTS Reduction

A LTS reduction is a procedure that is done on a plant every 3 to 4 years typically. It is a process that can have catastrophic consequences if not followed as instructed. Due to its infrequency, a thorough review of the procedure and potential consequences is imperative each time BEFORE performing the actual procedure. The CuO in the LTS catalyst is reduced to metal Cu by an exothermic reaction. To effect the reduction, hydrogen is carefully and precisely added to a carrier gas (natural gas or nitrogen) and the bed temperatures are closely monitored. Anything that can affect the hydrogen concentration at the inlet can lead to temperature run-away. Measures taken to achieve adequate control of hydrogen include: • complete isolation / blinding of any potential

sources of hydrogen other than reduction gas especially when at higher pressure than the carrier gas;

• gas composition analysis of the reduction and carrier gases. Natural gas carrier gas should be free of hydrogen recycle;

• reliable supply of reduction and carrier gas flows; and reduction gas flowmeter calibration at constant conditions of carrier gas flow and back pressure;

• Gas testing schedule of inlet and outlet hydrogen concentrations.



Caution statements in a procedure describe reducing or removing reduction gas at certain peak temperatures. Achieving optimum catalyst performance and preventing temperature run-away are both considered. Studies have been done to understand what happens at elevated temperatures that cause run-away. In the case of a natural gas carrier, CuO reacts with methane at 475oF (246oC) via the following reaction [5], CH4 + 4CuO → 4Cuo + CO2 + 2H2O (7) This is also exothermic creating even higher bed temperatures. At temperatures around 700°F (370oC), hydrocarbon cracking occurs which produces hydrogen which further reacts with the catalyst and leads to temperature run-away. Under such circumstances vessel design temperatures can be exceeded.

Methanator Operation

The operation of the methanator is seldom given much attention and COx concentrations are typically not measured. However changes in upstream conditions can affect its performance which can have an adverse effect on plant efficiency, production and even equipment. During start-up, the frontend of the plant is brought up and the synthesis gas is introduced into the synthesis loop. The inlet temperature of the methanator is typically controlled by available steam and the feed/effluent exchanger. Introducing make-up gas to the loop with lower than normal methanator inlet temperature can introduce higher levels of COx. These higher concentrations increase the potential for ammonium carbamate formation within the loop which fouls equipment and increases the loop pressure drop which in turn increases the energy requirements of the plant. In addition to ammonium carbamate formation, ammonia synthesis catalyst is temporarily poisoned by oxygenates which will adversely impact plant production and loop performance. Once the COx levels are returned to normal, production

rates can take quite some time before recovering. If ammonia synthesis catalyst is exposed to high COx levels for long periods of time, the poisoning affects can be permanent. In addition to operating conditions at the methanator, mal-operation or upsets at the upstream LTS reactor which cause increased inlet CO to the methanator can overload the available capacity of the catalyst and lead to higher COx leakage. Upsets at the CO2 removal unit tend to be more frequent and can have a greater impact on the methanator. CO2 is also more difficult to methanate than CO so higher levels can overwhelm the methanator resulting in premature COx leakage. In cases where the CO2 increases to such an extent that the resultant exotherm elevates the methanator temperature to equipment design limits, process gas should be immediately shut off or vented upstream. The operation and control of the CO2 removal system is key to maintaining the inlet levels to the methanator. In many plants that have been revamped or are running at higher than design throughputs, the CO2 removal system is pushed to run at maximum load. At these conditions, small fluctuations can lead to gross carryover of solution and CO2 breakthrough. Some solutions such as Sulfinol and Vetrocoke containing sulfur and arsenic respectively can poison and permanently deactivate the methanator catalyst, while others such as Catacarb and Benfield do not necessarily contain poisons but can leave potassium salt deposits on the top of the bed which reduces catalyst activity by blocking active sites and also leads to increased pressure drop. Severe liquid carryover of all solutions can also cause physical damage of the catalyst further increasing the pressure drop. A skim of the catalyst is usually performed to correct the situation.



Conclusion

In summary, understanding the basic principles of catalytic reactor operation in an ammonia plant can help to avoid costly performance and operating problems. Operations and technical crews should routinely review and analyze process monitoring systems and procedures to ensure that the consequences of mal-operation are understood and best available procedures and technology are in place. Understanding the impact on catalyst and process equipment not only from a safety perspective but also from a cost of operation standpoint is important so that potential adverse situations are avoided or minimized in magnitude. With this complete knowledge and understanding, high performance catalytic technology, such as offered by Sud-Chemie, can achieve its full utilization and maximize the efficiency and productivity of the plant.

References

[1] David Rice “The Value of Catalyst Performance in Ammonia Plants” AIChE 2002

[2] David Pearce, Venkat Pattabathula, Suresh Bhatia, Massimo Iob, Jim Richardson “Failure and Replacement of Ammonia Converter Baskets and Application of AmoMax Catalyst” AIChE 2008

[3] David Rice “Loading of Primary Reformer Catalyst Tubes” AIChE

[4] David Borzik, Ron Howerton, Taylor Archer “Successful Recoveries from Major WHB Failures Experience” AIChE 2007

[5] J. Richardson, J. Wagner, R. Drucker, H. Rajesh “Understanding Hydrocarbon Reactions During LTS Catalyst Reductions” AIChE 1996



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