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Methane steam reforming in laboratory scale
Farinha, J.F. [email protected]
Instituto Superior Técnico, Universidade Técnica de Lisboa
Abstract
This paper presents experimental results of methane steam reforming, at high temperatures, on
a laboratory scale reformer using a sample from a commercial catalyst based on Nickel. Several
parameters were changed, such as SCR (Steam Carbon Raio), temperature, SV (Space
Velocity) and reactor configuration to optimize the conditions for reforming, including the
conversion and selectivity.
SCR was changed from 1.5 to 4, with the lower values (below 2) showing carbon deposition,
while the best conversion was obtained at 2.5 and 3. The increase of SV led to a reduction of
conversion, although hydrogen generation always increased. The selectivity increased with
space velocity suggesting the catalyst includes elements favoring the formation of CO2.
The average reformer temperatures tested ranged from 690 ºC to 813 °C and SV since the
10000 h-1 to 50000 h-1. The temperature at the inlet was slightly higher than the furnace set
point while the outlet value was about 250ºC lower than the inlet. The increase of temperature
had an impact in increasing conversion and lowering selectivity as expected. The activation
energy of the apparent reaction rate was estimated as 95 kJ/mol in the range of literature
values.
Tests were performed using two different heights of the reactor (5 and 10 cm) showing a small
advantage of the later for a constant SV. Some tests were also carried out for a reactor in layers
of sand and catalyst with a total height of about 10 cm leading to the best conversion compared
with the other reactors. Using this procedure it is possible to obtain good results by reducing
the loading of catalyst by 50%
1 – Introdution Hydrogen has been listed as a strong candidate as energy carrier, contributing to solve the
current energy and CO2 problem. Recently a strong interest has arisen in using H2 based fuel
cells as future source of energy conversion due to the high conversion efficiency of hydrogen
energy to electricity in small scale as well as not having emissions of pollutant gases [1]. The
production of hydrogen in the long term may be derived from non CO2 emitting technologies,
such as electricity from renewable energy sources or from nuclear power.
In the shorter term the introduction of hydrogen relies on the use of the cheaper method to
produce it that is conversion from a fossil fuel. There are several processes or a combination of
up to five main reactions: four independent: Partial Oxidation (PO) catalytic or not (CPO-
Catalytic PO); Steam Reforming (SR) Water Gas Shift (WGS) reaction; PReferential OXidation
1
of CO (PROX) and a fifth: metanation [2]. PO has the advantage of being thermally self
sustained but it has a small yield of hydrogen. PO combined with steam injection and WGS
leads to a process called Auto-Thermal steam Rreforming (ATR).
There are several options regarding the fuel to produce hydrogen, the process being simpler
firstly for methanol owing to the lower temperature and then for hydrocarbons and in particular
for natural gas. in this case at high temperature (and pressure).The methane steam reforming is
the leading method of hydrogen production in the world [3], owing to the great volume of natural
gas, consisting mainly of methane, produced annually. Steam reforming is outlined in Eqs. (2)
and (3) [4]:
CH4 + H2O CO + 3H2, ΔH = +206 kJ/mol (1)
CO + H2O CO2 + H2, ΔH = -41.2 kJ/mol (2)
The mechanisms for steam reforming [5] consider that in parallel with reaction (2) the reforming
into carbon dioxide may also occur in parallel:
CH4 + 2 H2O CO2 + 4H2, ΔH = +165 kJ/mol (3)
In this case the water gas shift reaction (3) may occur in both directions.
The reforming process is highly endotermic and usually takes place in a tubular reactor at a
temperature in the range of 750 to 850ºC. For a nickel catalyst, one of the main problems is
coke formation that is avoided by increasing the steam carbon ratio. The control of temperature
in the reactor is another important factor and the removal of sulfur compounds, present for
instance in odorants used for natural gas.
Natural gas on an industrial scale is the most used method for the generation of hydrogen
today. The reformers consist of industrial furnaces where low quality fuels can be burned
generating heat to reformer tubes where the catalyst is. The tubes, are arranged in regions of
the furnace which can produce temperatures of operation close to ideal values around 750 °C.
In order to reduce the scale of the reformers for a decentralized production, several prototypes
and pre-series commercial reformers are currently being developed for productions lower than
50 kg/day. A range of applications can be identified for the application of compact reformers [2].
The purpose of this work is to test a laboratory scale test reactor and to analyze the influence of
the operating conditions and optimize them. The characterization is made for different operating
conditions, calculating the main parameters of reformation, such as selectivity and conversion.
Several test results are available in the literature regarding the influence of the operating
conditions. Lee et al [5] present the variation of conversion with temperature and SCR, showing
that the results follow thermodynamic equilibrium, possibly due to the low space velocity used
(5000 h-1). Results showing the influence of space velocity available from the supplier of the
catalyst used in this work are presented in Ventura [8] showing a decrease of methane
conversion increasing the space velocity from 30000 to 125000 h-1.
The aim of the present work is to test a reformer laboratory reactor and to evaluate and draw
conclusions about the influence of operating conditions, including the SCR, average reformer
temperature and space velocity in the main performance parameters of reforming (selectivity
2
and conversion). To compare the results with literature the activation energy of the apparent
surface reaction rate is also evaluated.
Following this introduction, section 2 presents the experimental test section and the procedures
used in the experiments. Section 3 presents the results and show the influence of the operating
conditions and the reactor configuration. Section 4 presents the main conclusions.
2 – Experimental reactor A schematic diagram of the SR experimental set-up used in this work is presented in Fig. 1.
This experimental installation is suitable to carry out thermal treatments with controlled gas
composition. To prepare the catalysts, the reactor can be used for drying or calcination and
during catalyst use it can be used to oxidize deposited carbon or to reduce the catalyst.
The reactor is suitable to test any of the reactions used in fuel processing below 1200ºC. For
auto-thermal reforming oxygen can be added in parallel with steam and methane. For methane
steam reforming a nickel based catalyst is used working at a temperature around 750 ºC and
pressure up to ten bar in the test section.
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Figure 1 - Schedule of the installation: 1 - Bomb; 2 - evaporator, 3 - Oven and tubular reactor inside, 4 -
Cooling system (including a condenser and heat exchanger), 5 - balance and deposit of condensed water,
6 to 9 - Controlling flow Gas: 6 - Oxygen, 7 - Natural gas or methane, 8 nitrogen, 9 Hydrogen, 10 -
Controller flow LPG.
The installation allows the control the flow rate of four gas through the flow controllers on the
lines indicated 6 to 9. Three of the flows are mixed upstream of the evaporator (2), while the
flow of oxygen (6) is mixed after the humidifier, before the entry of the reactor. LPG can also be
3
added through a specific inlet with a drain for excess liquid. To feed liquid fuels at working
pressure these can be mixed with water if compatible with the pump materials.
Water is supplied to the evaporator that operates at 200 °C and the steam is mixed with the gas
mixture (with the exception of oxygen). The mixture with steam may well reach the reactor
without intermediate condensation and then is heated in the oven depending on the operating
conditions. The oxygen is mixed only after the mixture is moistened to avoid the presence of
flammable mixtures in the evaporator or blender.
The reactor tube can be easily removed from inside the furnace where it is positioned through
plugs assembled under pressure. Two thermocouples can be inserted in the reactor tube, one
before and another after the catalyst bed to monitor temperature. The furnace temperature is
controlled by other thermocouples installed in the furnace connected to the control system.
The reaction products are dried by cooling them down to about 20ºC using a Peltier cooler and
the condensed water falls by gravity to a tank in a precision balance. This tank is emptied when
the level reaches a high value. Initially the pump circulates a large water flow to ensure that the
pipes connecting are full of water. The steam water that is dragged with the products can also
be detected in the gas analysis.
The operating pressure of the reactor is governed by a controller installed upstream of the gas
meter. This gas meter is used to monitor the gas products flow rate. There is a sample line
available to connect directly to a gas chromatograph and a main line for the products exhaust.
In the present work gas samples were taken every ten minutes during the tests in a syringe
through a membrane mounted in a T-connection to the main gas products exit line.
Most of the tests were preceded by a reduction treatment of the catalyst and when the previous
case had a small SCR an oxidation was also performed in advance. The furnace was then
heated up to the specified temperature keeping a small nitrogen flow through the reactor. The
experiments were then started by fixing the inlet flows and were carried out during 150 minutes.
For each test 15 gas samples were collected and injected in the gas chromatograph. The
manual procedure introduced some air leakage in the samples and these were detected in the
analysis and removed to normalize the gas composition from the reactor. When the total volume
detected by the gas chromatograph differed more than 10% from the syringe volume the
analysis was disregarded. The values presented in the experimental results are the average of
the analysis validated. Based on the measured gas composition two parameters were defined
that are the conversion (C) and selectivity (S) according to the following equations:
( )COCOCHCH XXXXC ++−=244
1 (4)
( )22 COCOCO XXXS += (5)
4
3 - Experimental Results A total of 25 tests were performed, most with a reactor height of 5 cm and SCR=3, while
nominal reactor temperature was changed from 800 to 900ºC and SV from 10000 to 50000. The
space velocity was defined based on the inlet flow at normal conditions and the volume of the
catalytic section of the reactor. Besides the height of 5 cm tests were carried out doubling this
height or by keeping the 5 cm of catalyst but putting the catalyst particles divided by
intermediate layers of inert (quartz sand).
Due to the endothermic reaction temperature in the reactor decreases and in general the inlet
temperature was about 30 ºC higher than the nominal and at the outlet about 220 ºC lower than
the nominal, that is there was a temperature reduction of 250 ºC. The average temperature is
therefore about 100 ºC lower than the nominal furnace temperature but as most reactions
should occur in the initial part of the reactor where temperature is higher the data is always
presented as a function of the nominal temperature. These differences in temperature are not
only a result of the reaction zone as well as effects from the furnace design. The reactor is
located in the second half of the furnace and hence there is some temperature gradient due to
end effects.
Figure 2 shows the conversion obtained in tests as a function of the SV, with the results
grouped by different values of SCR, nominal temperature and different constitutions of the fixed
bed reactor.
Figure 2 - Conversion according to SV, with the results grouped into different values of SCR, the oven
temperature and configuration of the reactor
Figure 2 shows as expected that the conversion increases with temperature, because the rate
of reformation of methane has activation energy of very high and therefore is highly affected by
temperature. The values of certain conversion rates for the smallest space velocities (and
longer times of residence) are closer to the conditions of equilibrium.
5
Changing the SCR from 3 to 2.5 has some influence but within the range of variability of the
results and therefore no firm conclusion can be made. Further tests performed with SCR from
1.5 to 4 keeping the inlet flow rate for methane constant and a constant temperature of 850ºC
showed lower values of conversion for the lower SCR which can be partly explained by the
deposition of carbon in the catalyst particles observed after these tests. For the higher SCR the
conversion also decreased but in this case this may be an influence of the larger space velocity.
Increasing the length of the reactor, keeping the space velocity constant led to a small
improvement of the conversion. This can be attributed to the larger area of heat transfer and an
increase of the heat transfer between the gas and bed, despite the single value of temperature
measured after the reactor being only slightly higher. If the comparison is made for similar inlet
flow rates of course the improvement is much more significant.
Using the reactor with several layers of sand and particles of catalyst, the conversion is higher
and has a smaller influence of SV changed from 24400 to 48900 h-1. The SV was defined based
on the volume of the catalyst, while if it was based on the overall volume SV would be from
9400 to 18800 h-1, showing still an improvement over the other cases. Although only two tests
were carried out for this configuration the results suggest a good improvement in conversion.
This is possibly due to the heating and temperature redistribution when the gases cross the
inert layers, leading to larger temperature at the entrance of the following catalyst layer in
accordance with [5]. In this configuration compared with the catalyst bed (H) of 10 cm it can be
observed that catalysts loads of less than about 50%, can achieve higher levels of conversion.
Figure 3 shows the selectivity as a function of SV, with the results grouped into different values
of SCR, nominal temperature and different constitutions of the fixed bed reactor, as above.
Figure 3 - Selectivity according to SV, with the results grouped into different values of SCR, the oven
temperature and configuration of the reactor
6
Figure 3 shows that in general the selectivity increases with SV and decreases increasing
temperature, the opposite behaviour of conversion. The effect of temperature can be explained
by thermodynamics that favours the formation of carbon monoxide at higher temperature,
lowering selectivity. Considering an average temperature it can be observed that the value of
selectivity observed in the experimental trials is larger than the equilibrium value. The catalyst
may contain some elements which encourage the formation of CO2 but when the residence time
is longer (smaller SV) this species can form CO, decreasing the selectivity.
Tests conducted at SCR = 2.5 show in figure 2 a slight decrease of selectivity, compared with
tests carried out on SCR = 3, due to the influence of the reaction of displacement of water.
Results obtained for a larger range of SCR keeping the methane flow rate constant show a
much larger increase of selectivity with SCR due to the larger water content as well as the larger
SV that is found to increase selectivity.
The selectivity has a small decrease in selectivity increasing the height of the reactor zone, that
is the opposite tendency observed for the conversion and similar behaviour is observed for the
layered reactor. This may be justified by the increase of the local temperatures.
Figure 4 shows the flow of hydrogen produced as a function of SV and the results were grouped
by nominal temperature of the reactor and SCR as above.
Figure 4 – Flow of hydrogen produced according to SV, with the results grouped into different values of SCR and temperature of the oven
From figure 4 it appears that the production of hydrogen always increases with the increase SV
and consequently with the reactants flow, despite the reduction of conversion because they are
associated with smaller residence times. It can also be observed in figure 4 that the hydrogen
production increases with temperature due to the effect of temperature in the kinetic rates. The
modification of SCR from 3 to 2.5 has a small impact in the production of hydrogen.
Interpretation of results with simplified models
Despite the results do not meet isothermal conditions and the composition of the catalyst is not
known in detail, the results were used to calculate reaction rate constants and the influence of
7
temperature is analysed., Estimative values of axial diffusion led to the conclusion that even for
the smaller flow rates the reactor is well represented by a plug flow reactor. Using this
simplification and assuming the reforming reactions to be first order in methane, allowed the
calculation of rate constants for the methane consumption (parallel in reactions 1 and 3). In
reality the rates also depend on the concentration of water vapour and the approach to
equilibrium conditions. Since there is always an excess of steam and the conversion is far from
complete, the use of a first order rate was considered a good approximation.
The rate constants determined were compared with mass transfer limitations outside the
particles, leading to the conclusion that the later has a small influence and therefore the rate
constants are characteristic of the apparent kinetics on the outer surface of the particles.
Figure 5 presents the calculated values as a function of the reciprocal of temperature. The
values obtained from the slopes allow to estimate the activation energy of 95 and 240 kJ / mol
for the two sets of values of SCR 3 and 2.5 respectively. Since most data was obtained for
SCR=3 the value of 95 kJ/mol is taken as more representative,, which is of the order of the
value obtained with mass transfer limitations within the particles (66 kJ/mol [6]). The value of the
true activation energy from the intrinsic reforming reactivity is larger 241 kJ/mol [7]..
Figure 5 - ln(K) as a function of the reciprocal of temperature (1/T).
4 – Conclusions
From this study it was found that the best oven temperature for operation is 850 °C, as the it
provides a good conversion and still high selectivity. Larger temperatures although may
increase conversion require more thermal input.
For the reactor with a fixed bed catalyst of 5 cm the best operating conditions are:
• Temperature of the furnace = 850 ° C;
8
• average temperature of reformation = 750 C (approximately);
• Space velocity = 42000 h-1 (approximately);
• SCR = 3.
The corresponding values for the conversion and selectivity of 78% and 42% respectively.
The influence of the operating conditions as expected show an improvement of conversion,
increasing temperature and reducing space velocity and in general the improvement of
conversion degrades selectivity.
Comparing the values of conversion and selectivity changing the reactor configuration it is
concluded that increasing that there is a marginal increase in conversion doubling the reactor
height, keeping the space velocity. The use of the catalyst in layers compared with the same
amount of catalyst in a continuous reactor shows a large improvement in conversion. Even
comparing similar volume reactors, one in layers partly filled with inert and the other only with
catalyst, the layered reactor increase slightly the conversion, using around 50% of the catalyst,
and continuing to produce high levels of selectivity.
The reaction rate shows a dependence in temperature that can be interpreted reasonably with a
first order kinetic rate, whose activation energy (95 kJ/mol) is found to be of the order of
literature values for apparent kinetic rates.
Nomenclature: SCR Steam Carbon Ratio
SV Space velocity [h-1]
T Temperature [K ]
K Rate of reaction [kg/m2s]
CPO Ctalytic partial oxidation
r Rate of reaction of methane [kg/m2s]
SR Steam reforming
hm Coefficient of mass transfer of [m /s]
WGS Water gas shift
PROX Preferential oxidation of CO
H height of the fixed bed of catalyst
9
References:
[1] Chen L., Hong Q., Lin J., Dautzenberg F.M., (2007), “Hydrogen production by coupled
catalytic partial oxidation and steam methane reforming at elevated pressure and temperature.”,
Journal of Power Sources 164 (2007) 803-808.
[2] Qi, A.; Peppley, B. e Karan K. (2007), “Integrates fuel processors for fuel cell application: A
review”, Fuel processing technology 88 (2007) 3-22.
[3] Dias J. A. C. e Assaf J. M. (2004), “The advantages of air addition on the methane steam
reforming over Ni/ʏ-Al2O3.”, Journal of Power Sources 137 (2004) 264-268.
[4] Fonseca A., Assaf E. M., (2004), “Production of the hydrogen by methane steam reforming
over nickel catalysts prepared from hydrotalcite precursors”, Journal of Power Sources 142
(2005) 154-159.
[5] S. Lee, J. Bae, S. Lim, J. Park, (2008), “Improved configuration of supported nickel catalysts
in a steam reformer for effective hydrogen production from methane”. Journal of Power Sources
180 (2008) 506-515.
[6] Akers (1955), W.W. and Camp, D.P. AIChE Journal, Vol. 1 (4): 471-475, “Kinetics of the Methane-steam Reaction”.
[7] Xu and Froment (1989) Methane Steam Reforming, Methanation and Water-Gas Shift: I. Intrinsic Kinetics, AIChE Journal 35 (1):88-96.
[8] Ventura, C. (2008), “Modelling of a reformer with integrated burner”, tese de mestrado.
Instituto Superior Técnico, pág. 6:8 e 71:72.
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