November 2021


Monitoring hydrogen plant performance—Part 2

Process monitoring is an indispensable practice to keep track of KPIs of an H2 plant.

Ramakumar, K. R., Johnson Matthey

Process monitoring is an indispensable practice to keep track of KPIs of an H2 plant. A good system of process monitoring not only ensures safe and reliable plant operations, but also helps operators to make strategic decisions, such as for catalyst changeout schedules. If KPIs are not monitored closely, there can be situations where the expected yields are not achieved. This affects the economics of the H2 plant and the entire refinery complex.

The main objectives of this article are to guide H2 plant process engineers in monitoring critical parameters and KPIs across each reactor in the H2 flowsheet, and in performing a detailed mass balance across the H2 flowsheet by using available information, such as dry analysis of outlet streams. Doing so will help identify bottlenecks across each reactor. The focus would be to see how much H2 is being produced before the final stream enters the pressure swing adsorption (PSA) unit.

This mass balance will also help estimate the outlet stream’s composition on a wet basis, thereby facilitating the estimation of equilibrium constants (Keq values) for steam methane reforming (SMR) and water-gas shift (WGS) reactions, which will help calculate the approach-to-equilibrium (ATE) values. These values, which are important in understanding the catalyst activity, can then be compared with the kinetic model values provided by the catalyst supplier.

Primary steam reformer

Every H2 process engineer knows that the primary steam reformer is the heart of the entire H2 flowsheet. The primary function of the unit is the conversion of methane and higher hydrocarbons (in the absence of a pre-reformer) in the feed to H2 (along with CO and CO2). Key parameters include the steam-to-carbon (S/C) ratio, outlet temperature, pressure, pressure drop, tube wall temperature (TWT), reformer firing and flame characteristics, and tube appearance. Methane conversion (as indicated by methane slip) and pressure drop are the two major performance indicators.

Approach to equilibrium is a theoretically important parameter in this section, as it indicates the performance of catalyst, while other parameters (such as outlet temperature, S/C ratio and pressure) are held constant. This value will be useful for the process engineers during the technical evaluation stage concerning catalyst offers. The reactions occurring in the reformer are equilibrium limited; therefore, the methane slip observed at zero (or close to zero) ATE indicates the minimum possible thermodynamically methane slip at the given conditions.

Another advantage of calculating the equilibrium temperature is that it provides an indication of the exact outlet process gas temperature. As the WGS reaction is quick, it can be assumed to always be in equilibrium at the reformer outlet conditions.

Normally, in most reformers, the reformer outlet temperature indication (TI) is located a few meters away from the tube outlet (where the catalyst ends). There is always some heat loss that needs to be assumed from the tube end to the TI point.

Mass balance across the primary reformer. The pre-reformer effluent is the feed to the reformer. Depending on the flowsheet, there could be some additional steam added at the reformer inlet. For reformers in hydrogen and carbon monoxide (HyCO) plants, recycled CO2 is normally added at the reformer inlet, with the intention to maximize CO yield. However, for the examples here, let us assume that no additional steam is added for Case 1 and that some steam is added for Case 2. Note: These are purely assumptions and bear no similarity to any existing unit.

Case 1. The total wet flow at the outlet of the pre-reformer—as per the calculation detailed in Part 1 of this article—was 4,028.4 kmol/hr. The reformer inlet temperature was 500°C, and the reformer’s outlet temperature was 840°C. Assuming 15°C heat less from the tube end to the TI, the tube outlet temperature would be 855°C and the outlet pressure would be 22 bara (21.7 atma). The outlet composition (dry basis) reported by the laboratory was the following:

  • H2 = 72.4%
  • CH4 = 4.57%
  • CO = 13.61%
  • CO2 = 8.1%
  • N2 = 1.35%.

In the following equation for this example, d is the outlet dry flowrate and s are the moles of steam at the outlet. The inlet and outlet moles in Case 1 are shown in TABLE 5. To calculate d using carbon balance, Eq. 20 is used:

907.47 + 62.16 + 0.373 = 0.0457d + 0.081d + 0.1361d                                                                                            (20)
d (the outlet dry flowrate) = 3,691 kmol/hr

To calculate s using O balance, Eq. 21 is used:

2 × 62.16 + 0.373 + 2785.32 = s + 2 × 0.081 × 3691 + 0.1361 × 3691                                                                     (21)
s (the outlet steam) = 1,809.73 kmol/hr

The outlet wet mol fraction is shown in TABLE 6. To calculate the SMR equilibrium constant, Eq. 22 is used:

Kp (SMR) = P2 × [([CO] [H2]3)/([CH4][H2O])] = 21.72 × (0.091 × 0.4863) / (0.031 × 0.329) = 482.3                            (22)

Eq. 22 can be substituted with Eq. 23 to obtain equilibrium temperature in °K:

Ln (1/Kp) = 0.2513Z4 – 0.3665Z3 – 0.58101Z2 + 27.1337Z – 3.2770                                                                           (23)

where Z = (1000/T)–1; T is in °K

Ln (1/Kp) = -6.179

Solving the equation using Excel, the answer is Teq = 846.3°C. Therefore, the approach to SMR equilibrium in the reformer is T – Teq = 855 – 846.3°C = 8.7°C. The WGS approach can be calculated using Eqs. 24 and 25:

Kp (WGS) = [([H2][CO2])/([H2O][CO])] = (0.486 × 0.054) / (0.329 × 0.091) = 0.8766                                                (24)

Ln (Kp WGS) = 0.63508Z3 – 0.29353Z2 + 4.1778Z + 0.31688                                                                                     (25)

Solving the equation, using Excel, Teq (WGS) = 845.9°C.

Case 2: The naphtha case. Assume that additional steam of 165 kmol/hr is added at the inlet of the reformer. The reformer’s inlet temperature is 500°C, and the reformer’s outlet temperature is 840°C. Assuming 15°C heat loss from the tube end to the TI, the tube outlet temperature is 855°C, and the outlet pressure is 23 bara (22.7 atma). The outlet composition (dry basis) reported by the laboratory was the following:

  • H2 = 68.53%
  • CH4 = 2.76%
  • CO = 15.86%
  • CO2 = 12.85%.

The outlet wet composition after completing C and O balances is shown in TABLE 7. Solving the equations as before, we receive the following:

Kp(SMR) = 538.3; Kp(WGS) = 0.864
Teq(SMR) = 851.4°C; Teq (WGS) = 850.1°C.

Therefore, the approach to SMR equilibrium in the reformer is T – Teq = 855 – 851.4°C = 3.6°C. The process engineer should ask the catalyst supplier what the expected approach is and compare the same with the calculated values.

Tube wall temperature. Besides catalyst activity, the approach also depends on operating conditions such as outlet temperature, pressure and composition, as well as how the reformer is being operated. Additionally, the approach and the methane slip depend on how the reformer is being operated. Improper heat distribution across the tubes will increase the approach; therefore, the methane slip would be more than expected, although the catalyst is perfectly normal in terms of its activity.

Tube wall temperatures need to be measured on a regular basis, using a simple handheld pyrometer or advanced thermal imaging equipment. The TWT of all tubes—after applying a suitable correction for background radiation—should indicate the reformer heat distribution balance. A spread (maximum–minimum) of more than 70°C normally indicates that there is scope for improving the heat distribution. The measured TWT of tubes can also be compared with the expected TWT profile as per the catalyst supplier’s kinetic model.

Pressure drop across the steam reformer. This is another critical parameter that needs close monitoring. The rate of pressure drop increase depends on how well the reformer is being operated. Due to the expansion and contraction of tubes, frequent trips or shutdowns can break the catalyst pellets. The broken pellets can contribute to a higher pressure drop. A sudden increase in pressure drop would need the time to be isolated and for any event in that period to be investigated. In most plants, the pressure drop is measured between the reformer inlet and waste heat boiler (WHB) or process gas boiler (PGB) outlet. Therefore, any fouling in the WHB/PGB can also affect the pressure drop. The process engineer should have a record of the number of trips/shutdowns and the reason for such a trip/shutdown. In many instances, a step increase in pressure drop is observed after starting up after a sudden trip.

Shift reactor

The primary function of the shift reactor is to convert CO formed in the reformer to H2 by reacting it with steam. The key parameters are inlet temperature, outlet temperature, steam-to-dry gas ratio, pressure drop and temperature profile. The KPIs include CO conversion, pressure drop and WGS approach.

Plotting the temperature profile clarifies how the catalyst activity declines with age. A clear trend can be obtained by plotting the percent exotherm profile rather than the temperature profile. The percent exotherm at any bed temperature equals [(T – Tmin) / (Tmax – Tmin)] × 100. A sample illustration of both bed temperature and exotherm percentage plots across a high-temperature shift (HTS) reactor is shown in FIG 3. Profiles 1 and 2 are different profiles for the same feed at different inlet temperatures. Not much can be inferred by plotting the bed temperatures. However, when the profiles are plotted by taking the exotherm percent at each position, it becomes clear that, by increasing the inlet temperature, the reaction profile becomes steeper.

FIG. 3. HTS bed temperature profile (top) and the percent exotherm profile across the HTS (bottom).

Pressure drop. The shift reactor, either HTS or medium-temperature shift (MTS), being downstream of the WHB/PGB, is vulnerable to fouling issues due to upstream boiler water leaks. This causes pressure drop to increase due to the buildup of boiler solids. The pressure drop trend should be carefully monitored, especially after a trip incident. In addition, any wetting incident can reduce the strength of catalyst pellets, which could contribute to a high pressure drop. Any inappropriate or inadequate hold-down layer on top can also cause high pressure drop issues.

Mass balance across the shift reactor. Assume there is an HTS reactor in the flowsheet and that mass balance will need to be calculated.

Case 1. The inlet temperature is 320°C, and the outlet temperature is 394°C. The outlet composition (dry basis) reported by the laboratory was the following:

  • H2 = 74.86 %
  • CH4 = 4.16%
  • CO = 3.38%
  • CO2 = 16.36%
  • N2 = 1.23%.

TABLE 8 shows the inlet and outlet compositions after doing C and O balances. At the end of the shift reactor, 3,038.25 kmol/hr of H2 is produced in Case 1. Using Eqs. 26 and 27, the WGS approach can be calculated as:

Kp (WGS) = [([H2][CO2])/([H2O][CO])]                                                                                          (26)
Ln (Kp WGS) = 0.63508Z3 – 0.29353Z2 + 4.1778Z + 0.31688                                                          (27)

Eqs. 26 and 27 are used to calculate the following:

  • Kp(WGS) = 10.2 and Teq (WGS) = 403.2°C
  • WGS approach = Teq (WGS)T = 403 – 394 = 9.2°C.

Case 2: The naphtha case. In this case, the inlet temperature is 320°C, and the outlet temperature is 396°C. The outlet composition (dry basis) reported by the laboratory was the following:

  • H2 = 71.96 %
  • CH4 = 2.46%
  • CO = 3.2%
  • CO2 = 22.37%.

TABLE 9 shows the inlet and outlet compositions after doing C and O balances. At the end of the shift reactor, 1,147.25 kmol/hr of H2 is produced in Case 2. The temperatures and WGS approach are the following:

  • Kp(WGS) = 10.8 and Teq (WGS) = 397.1°C
  • WGS approach = Teq (WGS)T = 397.1 – 396 = 1.1°C.

PSA section

Generally, H2 recovery across the PSA is 85%–90%. Assuming 87% H2 is recovered across the PSA, the H2 production in Case 1 would be 3,038.25 × 0.87 = 2,643.3 kmol/hr. In Case 2, it would be 1,147.25 × 0.87 = 998.1 kmol/hr. This needs to be cross-checked with the PSA purge gas flow and the H2 content.


There is no denying that proper checks and balances in the plant are crucial for ensuring maximum operational efficiency—and that mass balance is one of those crucial checks. The primary objective of this practice-oriented article is to ensure that the process engineer understands the significance of KPIs across each section of the plant and is confident in doing calculations (e.g., approaches to equilibrium), using available plant and lab information. By following the monitoring aspects highlighted in this article, the process engineer will be in a better position to make sound technical judgments when doing a technical bid and/or routine plant evaluations. HP


  1. Riazi, M. R., “Characterization and Properties of Petroleum Fractions,” ASTM International, West Conshohocken, Pennsylvania, 2005.
  2. Twigg, M. V., Catalyst Handbook, Second Edition, CRC Press, Boca Raton, Florida, 1996

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