How Hot is Your Reaction Furnace – Really? D’Arcy Blais Senior Projects Planning Advisor Imperial Oil darcy.h.blais@esso.ca Chuck Marshall Senior Projects Planning Advisor Imperial Oil chuck.w.marshall@esso.ca Dick Wissbaum Technology Director URS Corporation dick.wissbaum@urs.com

This paper discusses a real reaction furnace processing acid gas which contains a significant amount of ammonia in the combined acid gas feed. We will discuss design parameters – specifically, adiabatic temperature and residence time – in addition to predicted temperature and measured temperature. We will also present the results of

computational fluid dynamic modeling. We will present performance test data showing ammonia destruction in the reaction furnace. Finally, we will illustrate a method of

monitoring SRU pressure drop history. Introduction Experts in sulfur recovery have long agreed that the three keys to achieving satisfactory ammonia destruction are time, temperature and turbulence – the so-called “three Ts.” Time – the residence time in the reaction furnace – is easy to specify and to achieve. Turbulence is becoming better understood as burner vendors apply computational fluid dynamics to their burner designs. However, although experts readily agree on the required temperature – 2200 to 2280 °F (about 1200 to 1250 °C) – they often disagree on whether this is the measured temperature or the predicted temperature. This disagreement is made worse by the variety of methods for predicting temperature in the reaction furnace, and by the difficulty in measuring temperature in the reaction furnace. In this paper, we will explore this dilemma in the context of an operating sulfur recovery unit, and we will compare predicted operation with measured performance.

Design of the Sulfur Recovery Unit The sulfur recovery unit was designed to process the acid gas feed shown in Table 1. As in most refineries, there are two acid gas feeds: amine acid gas and sour water stripper acid gas. The combined acid gas feed contains 66% H2S, and 7% ammonia. The predicted adiabatic flame temperature in a single zone furnace was 2,220 °F (1215 °C). Based on discussions with the burner manufacturer and with the owner’s burner design expert, the decision was made to provide a single zone furnace, with the acid gas streams fed to dedicated nozzles in a single burner. Because the predicted flame temperature was at the low end of the acceptable range, it was decided to design for a 3 second residence time in the furnace. Operating Data without SWS Acid Gas By the time the sulfur recovery unit started up, the refinery crude slate had changed dramatically. The actual amine acid gas concentration – measured during an SRU performance test – is shown in Table 2. (Table 2 also shows the design basis concentration for comparison.) At initial SRU startup, the Sour Water Stripper was not yet running. Clearly, the increase in CO2 content at the expense of H2S content, and the resulting decrease in reaction furnace temperature were a cause of concern. Table 3 compares measured (by optical pyrometer) reaction furnace temperatures, with reported adiabatic and predicted adiabatic reaction furnace temperatures. Predicted reaction furnace temperatures were obtained using VMGSim 6.5, taking program defaults for all furnace parameters.1 Clearly, heat loss effects on temperature are significant in this reaction furnace. Reaction Furnace Heat Loss Traditional rules of thumb for reaction furnace heat loss are 3 to 5 percent of the furnace heat release, or 100 to 200 degrees F of temperature loss. The problem with these rules of thumb is that they reflect neither reactor size (i.e. unit capacity), physical design (i.e. refractory and thermal shroud design) nor turndown. In September 2011, Sikorski et. al. presented results of Computational Fluid Dynamic (CFD) modeling of heat transfer from a reaction furnace [reference 1]. Their model included the following effects:

1 The use of VMGSim in preparing this paper should not be construed to mean that the authors

consider other simulators to be less suitable. VMGSim was used because the program default settings provide reasonable results without recourse to empirical models. In this particular case, the acid gas compositions varied to such an extent that other software required different empirical models for different cases, and the temperature discontinuities introduced by switching models were of the same magnitude as the temperature discrepancies we were trying to resolve.

Radiation and convection from the bulk gas to the refractory hot face; Conduction through the refractory and shell; Radiation and convection from the shell to the air within the thermal shroud; Heat loss by flow of air through the thermal shroud; Radiation and convection from the outer surface of the thermal shroud.

The results indicated the following trends:

1. A very large change in process gas temperature (700 °F) resulted in a fairly small change in shell temperature (70 °F).

2. The temperature difference between the bulk gas and the refractory hot face was about 20 °F, and varied by only 5 °F from one end of the furnace to the other.

3. Calculated shell temperatures tended to fall at the lower end of the 300 – 650 °F desired temperature range with almost any thermal shroud design.

Unfortunately, they presented no quantitative data on the magnitude of the overall heat loss. URS uses two heat transfer models to estimate heat loss from the reaction furnace. The simple model assumes an outer shell temperature to estimate heat loss from the bulk gas to the refractory hot face and to the shell. The complex model – which is used to design the thermal shroud – adds heat loss from the shell to the air within the shroud, and heat loss from the outer surface of the shroud. The complex model is a multi-zone model to account for radial differences in temperature and heat transfer. Since we are not concerned with performance of the thermal shroud – and since the performance test data includes shell temperature measured at six points – the simple model is sufficient for this analysis. The simple heat loss model is shown in Figure 1, and the relevant equations are given in Table 4. The solution is found by trial and error – find a value of T2 such that QRAD + QCONV = QCOND. (Generally, T2 will be 20 to 50 degrees F below the bulk gas temperature.) Using this model, we can estimate reaction furnace heat loss – as a fraction of heat release – as a function of capacity and turndown. The results are shown in Figure 2.2 We can now reconcile the observed furnace temperature of 1860 to 1910 °F with the predicted adiabatic temperature of 2090 to 2220 °F – heat loss in this particular furnace at this particular turndown rate accounts for about 260 °F of temperature drop. It is important to recognize that, in a turbulent reactor, the bulk fluid temperature is the lower temperature – there is no appreciable temperature profile across (or along) the reactor.

2 Since most simulators now report enthalpy based on free energy of formation, it is not a simple

matter to obtain reactor heat release. Therefore, Figure 2 shows heat loss as a fraction of the enthalpy change from adiabatic flame temperature to 600 °F. This temperature has been chosen arbitrarily to represent the outlet temperature of the waste heat boiler, and is well above the dew point of sulfur so that heat effects from condensing sulfur do not confuse the issue.

This is comparable to a CSTR, in which the entire reactor volume contains reactants (and products) at their outlet concentrations. HEC Technologies – who supplied the reaction furnace and burner – prepared a CFD model of the furnace operating under Test “A” conditions. The results are presented in figures 3 and 4. (Figure 4 plots the same data using a smaller temperature span for better resolution of temperature distribution within the furnace.) The operating temperature at the furnace outlet predicted by CFD was 35 degrees below the measured operating temperature, as shown in Table 3. (The optical pyrometer is located about 12” away from the outlet nozzle and looks horizontally across the full diameter of the reaction furnace.) Operating Data with SWS Acid Gas It was clear based on both predicted and measured operating temperatures that ammonia destruction might not be adequate with the introduction of SWS acid gas. Predicted reaction furnace temperatures of about 2100 °F were 100 to 200 °F lower than desired for ammonia destruction. Therefore, the owner made plans to test the unit under a variety of conditions once sour water stripper acid gas was introduced into the unit. The following tests were conducted:

Furnace temperature and ammonia destruction with amine acid gas only; Furnace temperature and ammonia destruction with amine acid gas and fuel gas; Furnace temperature and ammonia destruction with amine acid gas, SWS acid

gas and fuel gas; Furnace temperature and ammonia destruction with amine acid gas and SWS

acid gas. This choice of test conditions allowed measurement of ammonia destruction under conditions of both low and high ammonia feed content, and low and high reaction furnace temperature. Figure 5 shows ammonia destruction – both as percent of ammonia in the feed, and as residual ammonia exiting the waste heat boiler – as a function of adiabatic reaction furnace temperature. The figure shows consistently high destruction as percent of ammonia in feed – greater than 99% - at temperatures above about 2380 °F. This is consistent with guidelines accepted within the industry. However, note that ammonia slip is well above 250 ppmv at 2390 and 2400 °F. Figure 6 shows ammonia destruction as a function of measured reaction furnace temperature. This figure shows acceptable destruction efficiency at a measured temperature of about 2125 °F – which is considerably below the industry guideline of 2200 – 2280 °F.3 3 Please note that the four data points on the left side of Figures 5 and 6 represent small

amounts of ammonia in the feed (generally less than 1 volume percent), so even modest ammonia destruction is sufficient to achieve acceptable residual ammonia levels.

In February 1999, Klint and Dale presented a significant amount of data relating ammonia destruction in commercial Claus units to furnace temperature [reference 2]. They chose to use a modified adiabatic flame temperature – that is, adiabatic flame temperature less a correction for heat loss – as the temperature of interest.4 They concluded that there is a minimum furnace temperature below which ammonia destruction cannot be achieved, regardless of residence time. They suggested a value of 2280 °F (1250 °C) for this minimum. Based on the test data from this particular furnace, it appears that temperatures as low as 2125 °F (1160 °C) might be acceptable. So what is the tolerable level of ammonia slip from the reaction furnace? Reference 2 suggests that operating with residual ammonia levels above 300 ppmv creates a significant risk of ammonia salt deposition in downstream equipment. However, it also notes that some facilities with measured residual ammonia levels above 300 ppmv operate without evidence of ammonia salt deposition. More recently, Alberta Sulphur Research Ltd. has demonstrated that ammonia salts can be deposited on cool tubes (such as condenser tubes) at ammonia concentrations as low as 60 ppmv [reference 3]. However, they suggest that the salt deposition rate might be low enough in a commercial facility that the condensed sulfur washes the salts out of the condenser and into the sulfur pit. This may provide an explanation of why some plants experience plugging at ammonia concentrations of 300 ppmv while others do not – differences in condenser design, cooling medium, liquid sulfur flow and the like may determine whether or not the ammonia salts, which are almost certainly formed in any unit, will accumulate in the condenser. SRU Pressure Drop Monitoring Based on the results of the performance tests summarized in the previous section, it was clear that the facility must either operate with fuel co-firing or accept higher levels of ammonia exiting the reaction furnace. Both of these strategies carry risk, especially since this facility has no natural gas supply, and the available refinery fuel gas is subject to significant composition changes. Given the potential for ultimate unit capacity reduction from fouling/plugging due to gas firing or ammonia salt formation, it was decided to closely monitor the SRU ∆P (pressure drop as measured by the pressure in the combustion air line) to ensure early detection of any unexplained increase in P. Since ∆P is proportional to the square of the flow through the unit, and flow is proportional to the individual stream feed rates, the fundamental equation is: ∆P = (C1·QAAG + C2·QSWSG + C3·QMAIN + C4·QTRIM + C5·QFUEL)2 4 We agree with this approach; however, an industry-standard heat loss model is desirable to

ensure a common basis.

Which can be rewritten as:

∆P0.5 = C1·QAAG + C2·QSWSG + C3·QMAIN + C4·QTRIM + C5·QFUEL The five constants, C1 through C5, were regressed against operating data during a “baseline” operating period. The resulting correlation was then used to predict “baseline” ∆P which was compared against measured ∆P. Figure 7 is a plot of the measured vs “baseline predicted” ∆P through the baseline period, and shows the high level of predictability that the regression provided through this specific period. Figure 8 shows a plot5 of the use of the “baseline correlation” to predict earlier periods of operation back to the initial startup of the unit. The plot clearly shows that for periods before the “baseline period”, the measured ∆P is lower than the ∆P predicted using the “baseline correlation”. This suggests that some form of gradual or step change plugging took place some time before the “baseline” period. To correct for this in the correlation, the “baseline correlation” was simply scaled to provide an alignment with start of run predicted vs measured ∆P. The resulting scaled correlation then represented the “clean ∆P correlation”. Lastly, by calculating and plotting a ratio of measured ∆P over “clean predicted” ∆P, a very useful plugging factor index was created to monitor trends in ∆P tendency and to identify specific past ∆P tendency step change events. Figure 9 shows a plot of the measured vs “clean predicted” ∆P since the initial unit startup, as well as the calculated plugging factor index. Figure 10 shows a plot utilizing the “clean ∆P correlation” in conjunction with supporting operating data to identify the causes of the observed ∆P step changes. Analysis shows that since startup, the ∆P tendency has increased primarily due to three distinct and lasting ∆P tendency events, and to a lesser extent due to gradual fouling processes. Comparison with supporting operating data suggests that the first two ∆P tendency step increases followed periods of hot standby operation. Although specific testing has not been performed to determine the direct cause of the increased ∆P, the most likely explanation is soot formation and deposition on the catalyst beds due to hot standby operation fuel gas firing with insufficient combustion air. The third event in early October followed a brief outage of acid gas feed. Aside from these events, there is some evidence of gradual ∆P increase, however the impacts are minor compared with the potentially preventable step change events. Overall, we found the use of this monitoring technique to be very useful. The technique can readily be adopted at any point in time, and does not require that the “baseline” period be established at the start of run. Once a tool such as this is developed for a refinery, it can be used to monitor measured versus predicted clean ∆P to assess ongoing operations, or to study earlier periods to identify and learn from events that have resulted in significant ∆P increases.

5 The “baseline period” in Figures 8, 9 and 10 included periods with and without SWS acid gas

feed to the SRU. This was done to ensure that the resulting correlation represented both modes of operation. Operation with SWS acid gas was limited to the latter weeks shown in the figures.

Summary First, it is important to note that we do not advocate nor recommend operating any sulfur recovery unit at temperatures or residual ammonia levels comparable to the values shown in any of the tables or figures. Every sulfur plant is unique, and Reference 2 notes that some facilities experience plugging by ammonia salts at ammonia concentrations below 300 ppmv, while other facilities experience no plugging when operating at higher concentrations. Second, it should be clear that measured reaction furnace temperatures often operate several hundred degrees cooler than predicted temperatures. This may not be due to poor predictions by the simulator nor to improper temperature instrument calibration. This temperature difference depends not only on the mechanical configuration of the reaction furnace, its refractory, and the thermal shroud, but also on process throughput. Third, the answer to the question “which is the important temperature to achieve desired ammonia destruction?” is clearly: actual operating temperature. Unfortunately, this answer is unsatisfying for the following reasons: (1) the designer has no prior access to the actual operating temperature; (2) the actual operating temperature may be subject to fairly large measurement errors; and (3) the industry-accepted guidelines for required temperature likely evolved before reliable measurements were available, and therefore could be conservatively high. Therefore, the industry needs a common definition of, and means to calculate, expected operating temperature. We suggest that expected operating temperature should be based on predicted adiabatic temperature and corrected for heat loss from the furnace, at the operating conditions of interest. Fourth, a simple heat loss model, which is suitable for implementation using standard spreadsheet software, has been presented. Comparisons of expected operating temperature using this model and VMGSim with observed operating temperature have been provided. This model, or a comparable model, should be used by the designer to ensure that the actual furnace operating temperature will be acceptable. Finally, an empirical model to predict pressure drop in a given unit has been presented, and it has been shown to be effective both for long-term pressure drop monitoring, and for identifying and diagnosing upset conditions which led to a sudden increase in pressure drop. This may be of assistance to plants which are dealing with plugging issues, whether from ammonia salts or from other causes. Acknowledgements The authors wish to express their appreciation to Imperial Oil Limited for their gracious permission to present these results. In addition, special thanks are due to Nick Roussakis of HEC Technologies for his invaluable assistance in preparing the CFD results.

References

1. Sikorski, D., Huffmaster, M., Roussakis, N. and Corriveau, A., ”Reaction Furnace Weather Shield Design”, 2011 Brimstone Sulfur Recovery Symposium, September 13 – 16, 2011, Vail, CO.

2. Klint, B. and Dale, P., “Ammonia Destruction in Claus Sulphur Recovery

Units”, 49th Annual Laurance Reid Gas Conditioning Conference, February, 1999.

3. Clark, P., Dowling, N., Huang, M., Bernard, F. and Lesage, K., “Deposition of

Ammonium Salts in Claus Systems: Theoretical and Practical Considerations”, 2006 Brimstone Sulfur Recovery Symposium, September 11 – 15, 2006.

Table 1

Acid Gas Feeds to SRU (mole percent)

Component Amine Acid GasSWS Acid

GasCombined

Acid Gas H2O 4.544 32.370 10.697 Nitrogen 0.005 0.126 0.032 CO2 19.002 2.674 15.392 H2S 75.566 31.675 65.861 Ammonia 0.129 31.377 7.038 Methane 0.081 0.022 0.068 Ethylene - 0.046 0.010 Ethane 0.096 0.024 0.080 Propane 0.367 0.169 0.323 i-Butane 0.008 0.009 0.008 n-Butane 0.119 0.199 0.137 i-Pentane 0.007 0.017 0.009 n-Pentane 0.007 0.043 0.015 C6+ 0.069 0.056 0.066 HCN - 1.192 0.264 TOTAL 100.000 99.999 100.000

Table 2

Amine Acid Gas Composition (mole percent)

Component SRU Design BasisPerformance

Test #2Performance

Test #3 H2O 4.544 3.959 14.690

Nitrogen 0.005 .046 0.289

CO2 19.002 28.630 27.827

H2S 75.566 65.360 55.837

Ammonia 0.129 0.371 0.000

Methane 0.081 0.187 0.157

Ethylene - - -

Ethane 0.096 0.250 0.206

Propane 0.367 0.735 0.618

i-Butane 0.008 - -

n-Butane 0.119 0.291 0.241

i-Pentane 0.007 - -

n-Pentane 0.007 0.044 0.029

C6+ 0.069 0.071 0.061

BTEX - 0.053 0.041

COS - 0.003 0.003

TOTAL 100.000 100.000 100.000

Table 3 Reaction Furnace Temperature

(Degrees F) Test A Test B Measured 1,908 1,860 Reported Adiabatic 2,154 2,079 Predicted Adiabatic (VMGSim v. 6.5) 2,217 2,094 Predicted w/Simple Heat Loss Model 1,946 1,842 Predicted by CFD Model 1,873 ---

Table 4

Equations for Simple Heat Loss Model QCONV = hF·(·D1·L)·(T1 - T2) NRE = D1·V·/ V = W / / (·(D1/2)2) hF = 0.023·(kG/D1)·NRE0.8·(CP·/KG)0.3 QRAD = 0.173·e·(·D1·L)·[((T1 + 460)/100)4 - ((T2+460)/100)4] QBRICK = 2··kB/ln(D2/D1)·(T2 - T3)·L QREF = 2··kR/ln(D3/D2)·(T3 - T4)·L QCOND = QBRICK = QREF QCOND = XB·(T2 - T4)/(1+XB/XR) XB = 2··KB·L/ln(D2/D1) XR = 2··KR·L/ln(D3/D2) Symbol Definition Units CP Specific heat of process fluid Btu/Lb·F D1 Inside Diameter of Furnace / Hot Face Brick ft D2 Inside Diameter of Refractory ft D3 Inside Diameter of Vessel Shell ft e Emissivity of Process Gas --- hF Film heat transfer coefficient for convection Btu/hr·ft2·F KB Mean Thermal Conductivity of Hot Face Brick Btu/hr·ft·F KG Thermal conductivity of process fluid Btu/hr·ft·F KR Mean Thermal Conductivity of Refractory Btu/hr·ft·F L Overall Length of Furnace ft NRE Reynolds Number --- QBRICK Heat transferred through the Brick by conduction Btu/hr QCONV Heat transferred to the Brick by Convection Btu/hr QCOND Heat transferred through Brick/Refractory by conduction Btu/hr QRAD Heat transferred to the Brick by Radiation Btu/hr QREF Heat transferred through the refractory by conduction Btu/hr T1 Process Gas Temperature F T2 Temperature at Hot Brick Face F T3 Temperature at Hot Face Brick/Refractory Interface F T4 Temperature of Shell F V Velocity of process fluid ft/hr W Process Fluid Mass Flow Rate Lb/hr Viscosity of bulk process fluid Lb/ft·hr Density of process fluid Lb/ft3

Figure 1 Simple Heat Loss Model (see Table 4 for equations)

Reactor Shell, ID=D3

Insulating Refractory, ID=D2

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Figure 2 Reaction Furnace Heat Loss

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Figure 7 “Baseline Predicted” vs Measured SRU ∆P (Baseline Period Only)

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Figure 8 “Baseline Predicted” vs Measured SRU ∆P (Full Period from Startup)

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Figure 10 ∆P History Showing Step Change Events

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