primary reformer incident that ends in catalyst tube failure · production plants from waste heat...

Primary Reformer Incident That Ends in Catalyst Tube Failure

During a site-wide steam upset, an ammonia plant at Agrium’s Redwater Fertilizer Operations experienced catastrophic failures of four catalyst tubes in the primary reformer. The incident

investigation revealed several human factors and metallurgical issues that resulted in the failures. The post-incident mechanical integrity inspections and metallurgical analysis also revealed several

issues related to hot piping components that were not related to the incident.

Charles Ormsbee, David Craig, Alan Roe, John Mason Agrium Partnership

Introduction

n July 18, 2012, Agrium experienced the catastrophic failure of four catalyst tubes in the primary reformer of the Redwater (Alberta, Canada) Ammo-

nia 1 facility. The initiating event was a site-wide steam system upset that began over 3 hours before.

Background

The Agrium Redwater site is a highly integrated fertilizer facility. It consists of two separate op-erating units – a nitrogen fertilizer unit and a phosphate fertilizer unit – with shared utilities. Steam is the primary shared utility, which is ex-ported and imported from various production plants at three pressure levels – 600 psig (≈4100 kPag), 185 psig (≈1300 kPag) and 30 psig (≈200 kPag) - depending on their needs. The utilities unit makes up the difference to keep the site in balance.

The Agrium Redwater Ammonia 1 plant was commissioned in 1968 and has an original nameplate capacity of 600 stpd (544 MTPD). It was debottlenecked in 1998 to its current nomi-nal capacity of 840 stpd (762 MTPD). For the past 15 years the plant has been operated in block production mode depending on the site ammonia balance and the local demand for an-hydrous ammonia. Due to the block operation, the catalyst tubes, which were replaced in 1997, had approximately 48,000 hours of operation at the time of the incident. The design of the plant is that of a conventional ammonia process with a TEWC (totally en-closed water cooled) electric motor for the syn-thesis gas compressor. The primary reformer is a Foster Wheeler™ terrace walled reformer with two cells. Each cell has a single row of 80 non-staggered catalyst tubes. With the installation of the electric motor, the plant can either export or import 600 psig steam to help balance the site steam balance by adjusting the firing on the aux-iliary burners in the primary reformer.

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612014 AMMONIA TECHNICAL MANUAL

Site Steam Balance

The site 600 psig steam is the pressure level of steam that is used for process steam in the Am-monia 1 plant. In addition to Ammonia 1, 600 psig steam is produced in the following production plants from waste heat and exported to the site:

Sulfuric Acid Plants (SA 1 and SA 2) Ammonia 2 (Plant 09) Nitric Acid The utility boilers make up the difference in production and demand from three natural gas fired boilers – two 50 ton/hour (45 MT/hour) and one 100 ton/hour (90 MT/hour). During normal operation, the utilities unit exports ap-

proximately 20 ton/hour (18 MT/hour) of 600 psig (≈4100 kPag) steam. The major importers of 600 psig (≈4100 kPag) steam are as follows:

Urea Utilities – to produce 185 psig and 30 psig

from extraction turbines for use in various production units

Ammonia Plant Cooling Water Systems (Plant 32 and Plant 36)

Letdown – to make up the difference in 185 psig and 30 psig supply and demand

The Redwater site steam balance is shown in Figure 1. Ammonia 1 is indicated as Plant 01 in this figure.

Figure 1. Redwater site typical steam balance (all pressures, #, are psig; all flows are thousand lb/h)

62 2014AMMONIA TECHNICAL MANUAL

Anatomy of the Incident

Site Wide Events

The incident began at approximately 09:00 on July 18, 2012, when Sulfuric Acid 1 plant (SA 1), circled in blue in Figure 1, was shut down for unplanned maintenance. Within four minutes, ≈40 ton/hour of 600 psig steam pro-duction was removed from the site (the differ-ence between the 75 ton/hour of waste steam production and 35 ton/hour of steam consump-tion by the SA 1 air blower circled in blue). As shown in Figure 2, the site 600 psig steam head-er pressure begins to decline and the production from utility boilers 1 and 2 increases (utility boiler 3 was down for routine annual mainte-

nance and inspection). Within 5 minutes the two utility boilers are firing at maximum capaci-ty and producing an incremental 39 tons/hour of 600 psig steam exported to the site. The steam header pressure continues to decline because the 35 ton/hour of 600 psig steam consumed in the SA 1 air blower is exported as 30 psig steam to the phosphoric acid plant. As a result, the actual amount of 600 psig steam removed from the system with the shutdown of SA 1 was 75 ton/hour as the automatic letdown stations, circled in red in Figure 1, opened to make up the difference in the 30 psig steam demand. In hindsight, this situation was clearly unsustaina-ble without steam shedding.

Figure 2. Site 600 psig steam system parameters during incident

At approximately 09:10 utility boiler 2 tripped and the steam header pressure continued to de-cline. The root cause failure analysis (RCFA) concluded that the boiler tripped on flame out after the combustion air forced draft fan (com-bustion air) tripped due to a faulty low lube oil pressure switch. At approximately 09:20 the site initiated its steam shedding emergency procedure. The steam shedding procedure was completed at 09:30 with the shutdown of the Urea plant. The steam system was recovering until the Ammo-

nia 2 plant tripped at ≈9:45 and stopped export-ing steam.

Ammonia 1 Events

At the time of the incident Ammonia 1 was im-porting a small amount of 600 psig steam, there-fore it was dependent on the site steam header pressure to maintain the flow of process steam to the primary reformer. As shown in Figure 3, operations maintained ammonia production until it was apparent that the site steam system was not going to recover quickly. At approximately

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Site 600# Steam

System Pressure, psig

Steam

Production, thousand lb/hour

Time, July 18, 2012

Steam Production BLR No. 1

Steam Production BLR No. 2

Utility 600# Steam Export

Utility 600# Steam Pressure

SA 1 s/d

Boi ler 2

trip

Steam Shedding

Urea s/d

Ammonia 2 trip


10:10 the process steam flow to each cell dropped dramatically. Operations responded by tripping the backend, reducing the feed gas rate to 35% and decreasing the front-end pressure by opening the process vent downstream of the high temperature shift (HTS vent). In addition,

at 10:20, the process air compressor was tripped to reduce the power steam consumption. It had been left on line in an attempt to continue to produce steam in the secondary reformer waste heat boiler.

Figure 3. Ammonia 1 process parameters during incident

The steam flow stabilized, but inexplicably the HTS vent pressure was increased and the steam flow dropped again. At this point, ≈10:30, the feed gas flow was tripped and the reformer ter-race burner fuel gas was manually tripped. The fuel gas to the bottom burners was reduced, but was not stopped, to keep the reformer warm in preparation for a quick restart. Within 7 to 8 minutes the process steam flow went to zero and stayed there for approximately 20 minutes. At 10:55 a small flow of process steam was reestablished, but it is unknown whether this flow is due to a failed tube or that it was flowing to the HTS vent. There is some ev-idence that the flow may have been due to a failed tube based on the pressure of the steam header and the HTS vent, and a sudden change in the primary reformer firebox pressure. In any event, at some point between 10:55 and 11:10 four catalyst tubes failed (primary reformer fire-box pressure readings indicate the failures were likely at different times).

At some point before 11:00, the exact time is unknown, the field operator opened the manual emergency steam valves to cell 1 and cell 2 of the primary reformer, F-101. This steam is used to keep the feed gas coils in the convection sec-tion of F-101 cool during startup when there is low or no feed gas flow. Field operators report hearing a ‘pop’ sometime between 11:00 and 11:10. Within a few minutes they discovered the failed catalyst tubes in F-101. The four failed tubes were all in cell 2. The fuel gas to the reformer was stopped by 11:15.

Root Cause Failure Analysis

Operational

The focus of this paper is on the root cause fail-ure of the primary reformer catalyst tubes, not the site-wide steam upset.

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Pressure, psig

Steam

Flow, thousand lb/hour

Feed Gas Rate, %

Time, July 18, 2012

F‐101 Steam Flow Cell 1

F‐101 Steam Flow Cell 2

F‐101 Feed Gas Rate

NH3 1 600# Steam Pressure

HTS Vent Pressure

Feed gastrip

Process Steam

reestablished

Front‐end Pressure

reduced


While the cause of the failures may seem obvi-ous – the tubes were overheated and failed due to continued heat input with insufficient flow – it is not that simple. As with most serious inci-dents, the superficial cause often does not get to the real causes(s) that, if eliminated, will pre-vent the incident. As a result, we often make changes that only prevent the superficial cause and the incident is doomed to be repeated. During the event described above, a second is-sue was taking place. The operator’s attention was almost entirely on the steam upset, and the operation and condition of the reformer. To maintain this focus, the operator had tripped the back-end of the plant and the process air com-pressor. The mindset at the time was that the event was the same as a start-up – the reformer was warm, on steam, and ready for gas introduc-tion. Unfortunately, there are some subtle dif-ferences between an actual plant startup and the condition the plant was in at the time of the

event. These differences, and operations failure to realize that they were following a routine, or normal, procedure (start-up procedure) versus an emergency procedure, eventually resulted in four failed catalyst tubes. When the plant is started from cold following the normal start-up procedure, the front-end waste heat boiler steam drum, D-105 in Fig-ure 4, level is controlled like any other boiler, via a level control valve (LV-005 in Figure 4). The boiler feed water (BFW) via LV-005 flows through the economizer coil in the convection section of the reformer, F-101 in Figure 4. As the startup progresses BFW is eventually re-quired to cool the inlet gas to the low tempera-ture shift (LTS) in a second economizer, E-103 in Figure 4, via FV-007. When the back-end is brought online, additional BFW is introduced to the steam drum via a third economizer, E-138 in Figure 4, via FV-008.

Figure 4. Ammonia 1 process parameters during incident


During this event, with all the alarms and activi-ty in the control room, the operator knew that the level control, LV-005, was in automatic mode and assumed it would maintain level (as a level control is apt to do). The level control did close, but FV-007 and FV-008 were also in au-tomatic mode and continued to feed BFW into D-105 at a combined flow of approximately 400 USGPM (≈90 m3/hr). Post event calculations confirmed that D-105 and T-101, the steam wash column, were com-pletely filled with water at approximately the time that the catalyst tubes failed. No evidence of water was found in any of the other 600# steam consumers, however, the emergency steam lines are the only connection that is on the bottom of the steam header. It is surmised that BFW flowed by gravity from the bottom of the steam header to the mixing tee at the inlet head-er to the catalyst tubes. The first take off for the emergency steam is to cell 2, so the water flowed to it first. When the process steam flow was reestablished, the steam carried the water into the catalyst tubes and resulted in a steam explosion. While not overly severe as others1,2 it is eerily similar to lesser events that have been reported3. The fact that only four tubes failed is not fully understood, but it is a reasonable pos-sibility that the four failed tubes were mechani-cally weaker than the others due to being hotter (clearly there was no effort to balance the firing during the event and some tubes were likely hot-ter than others) or due to more advanced in-service creep damage. On the human factor side, the emergency oper-ating procedure for loss of steam does not clear-ly identify who has the responsibility to make the decision on the action required and who is responsible for carrying out that action. It is tacitly understood that the responsibility is the panel operator. This understanding is informal-ly (verbally) passed on to the panel operator during training.

Compounding the decision making of the panel operator is that the site had experienced two similar events in the previous two weeks, and in both events the ammonia plant was kept on-line. The belief of everyone, including senior person-nel, was that this was a recoverable event – i.e., normalization of deviation.

Furnace Inspection and Repair

The post incident inspection activities included inspection of the inlet header, inlet pigtails, re-former catalyst tubes, outlet header and outlet pigtails.

Figure 5. F 101 Primary reformer

Inlet Headers

The inlet headers are 10 inch diameter Schedule 80 pipe, 304 H material and each header is 45 feet (~13.7 m) long. The inlet header inspection found a 4 ½ inch upward bend from the center of the header to the outer ends. This damage was not the result of the incident or any previ-ous incident but was the result of the stress re-laxation during operation. The inlet headers were changed in 1998 from 2 ¼ Cr to 304H stainless steel and in fabrication no heat treat-ment was performed after the pigtail couplings were welded to the header. The header then took an in service bend due to stress relaxation from the operating temperature.


New headers were manufactured from 304H material and heated treated to relieve the weld-ing stresses. To date these headers have re-mained in the straight position.

Figure 6. Inlet Headers Removed from Furnace

Inlet Pigtails

The inlet pigtails are 1 ¼ inch schedule 80 pipe, 304 H material and a Z configuration with hori-zontal connection to the catalyst tube and a top connection to the inlet headers. Inspection of the inlet pigtails found cracking at the both the catalyst tube connection and at the header con-nection. In cell #1 the inspection found 18 cracked pigtails, 16 at the connection to the header and 2 pigtails at the catalyst tube connec-tion. In cell #2 the inspection found 13 cracked pigtails, 5 at the top connection the header and 8 at the bottom connection to the catalyst tube. Metallurgical examination of the failed pigtails determined the cracking was from stress relaxa-tion cracking. There was no evidence to suggest any of the materials were defective.

Figure 7. Cracked Pigtail #44 on East Header The pigtails were replaced in 1998 along with the inlet header and the material also was changed from 2 ¼ Cr- 1 Mo to 304H SS. After finding this amount of cracking in the pig-tails and the bend in the inlet headers it led to further investigation on the tube support hang-ers, header supports and inlet piping supports. Several piping flexibility cases were examined and the final conclusion is that the upward bend in the header contributed to high stress levels at the pigtail connections. The highest stress levels in the pigtails occurred at the outer ends of the header since this loca-tion has the greatest movement of the header (from bending). This movement resulted in a large bending moment at the pigtail to header connection. The combination of high bending stresses and operating in the temperature range of over 1100°F (593°C) resulted in stress relaxa-tion cracking of the pigtails. New pigtails were manufactured and installed with the new header.


Figure 8. Inlet Pigtail Assemblies

Furnace Catalyst Tubes

The catalyst tubes are 4.057” (103,5mm) OD by 0.347” (8.81mm) MW and 47 feet (~14.3 m) long, Centralloy 4852 Micro material. The overheating of the catalyst tubes during the in-cident led to the rupture of four tubes (#17, #23, #27 and #71), all in Cell #2 of the Primary Re-former. The inspection of the catalyst tubes consisted of examination of two catalyst tubes to determine the cracking mechanism at the rup-ture areas. The metallurgical analysis of the catalyst tubes found that the tubes failed by creep. It also con-cluded that the creep damage was localized to the failure area and the pressures and tempera-tures reported during the incident were not high enough to result in tube rupture. A temperature close the material melting range of 2300 °F to 2500 °F (1260 to 1371 °C) was required to cause the tube rupture. The analysis also con-cluded that the creep damage displayed required higher than reported operating tube metal tem-peratures to result in the level of creep damage observed, given the service life of the tubes. It was also reported that the top and bottom por-tions of the tube did not display the same amount of creep damage as the failure area. The likely scenario is that these tubes overheat-ed substantially during the incident due to the loss of process flow through the tubes. All cata-

lyst tubes were replaced during the furnace re-pair. The time from the incident occurrence to completion of repairs was about 3 months.

Figure 9. Ruptured catalyst Tube, Cell #2

Outlet header

The outlet header is alloy 800 H material and is constructed with two 8 inch diameter headers connected to the catalyst tubes by 1 ¼ inch (31,75 mm) diameter pigtails, the headers are jointed together by a 10 inch collection header that tees to a transfer line leading to the second-ary reformer. The header operated since 1971, a total of 270,000 hours. Dye penetrant inspection was performed on all circumferential welds on the headers and at-tachment welds on valves connected to the header. The inspection revealed nine crack indi-cations on the west header. On the east header inspection, eleven in-situ metallography replicas were completed; one plug sample from base material and one boat sample from a weld were removed from the header. The plug sample and boat sample were taken to determine the depth of cracks observed in the in-situ replicas and to confirm the depos-ited weld metal chemistry. The eleven in-situ metallography replicas all showed partially sensitized parent metal, grain sizes within the specified range for Alloy 800H,


there were no creep voids and deposited weld metal microstructures were consistent with nickel alloy welded structures. The plug sample inspection found 200 μm deep inter-granular cracks. The boat sample inspec-tion had chemistry consistent with Inco-Weld A and there were aligned creep voids in the ID, OD and midwall sections. In summary, the plug and boat sample cracks were the result of creep and the cracks observed at the in-situ replicated areas can also be at-tributed to creep damage. The welds in all the assemblies displayed creep damage. The entire outlet header assembly was replaced with new material from storehouse inventory.

Figure 10. Outlet Header and Pigtails

Outlet pigtails

The outlet pigtails are 1 ¼ inch schedule 80 pipe, Alloy 800 H material and the incident re-sulted in a failure on #17 pigtail. There were three outlet pigtail samples (#17, #18 and #23) from Cell #2 that were submitted for metallurgi-cal examination. Pigtail #17 had an extensive circumferential crack at the catalyst tube end and the crack dis-played a “thick lip” appearance. The crack fea-tures and the surrounding area were consistent with creep damage. The size of this crack (about 50% of circumference) would result in a

significant process leak during operation and the conclusion is this crack was opened during the incident since catalyst tube #17 was one of the ruptured tubes. It is likely that movement of the tube after rupture provided the additional load to crack the pigtail. Examination of the other two pigtails displayed no obvious damage during visual, dye penetrant and metallographic inspection. All pigtails were replaced during the replacement of the header assembly.

Figure 11. Cracked Outlet Pigtail

Summary

In summary, the root causes that resulted in four primary reformer catalyst tubes catastrophically failing are as follows:

Heating the tubes above a safe operating temperature in the absence of flow;

Introduction of wet steam to hot tubes caus-ing a steam explosion;

Lack of clear and specific lines of responsi-bility documented in emergency procedures (i.e., who, when, what and how); and

Normalization of deviation.

Conclusions

The lessons learned from this event are as fol-lows:


Operating procedures, including emergency operating procedures, must clearly docu-ment who is responsible for each action and when those actions shall be carried out.

The review of incidents must include a re-view of the potential for near misses to nor-malize deviations, and, if it appears normali-zation of deviation is occurring, conduct refresher training (even if informally).

The Inspection, Testing and Preventive Maintenance (ITPM) of all protection sys-tems must be viewed as a high priority and not a nice-to-have that can be deferred for expediency.

Piping components such as headers con-structed with many branch connections and intended for high temperature service re-quire heat treatment to relieve welding stresses.

References

1 Rogers, M. “Lessons Learned from an Unusual Hy-

drogen Reformer Furnace Failure”, Safety in Ammo-nia Plants and Related Facilities Symposium, AIChE Technical Manual, Vol. 47, 2006.

2 Rani, B. “Catastrophic Failure of Reformer Tubes at Courtright Ammonia Plant”, Safety in Ammonia Plants and Related Facilities Symposium, AIChE Technical Manual, Vol. 47, 2006.

3 Taylor, W. et al. “Steam Explosion in an Ammonia Plant”, Safety in Ammonia Plants and Related Facili-ties Symposium, AIChE Technical Manual, Vol. 45, 2004.


primary reformer incident that ends in catalyst tube failure · production plants from waste heat...

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