Performance of a 100 kW Heatpipe Reformer Operating on Lignite

Article pubs.acs.org/EF

Performance of a 100 kW Heatpipe Reformer Operating on Lignite Jonas M. Leimert,* Peter Treiber, Michael Neubert, Aaron Sieber, and Jürgen Karl Chair of Energy Process Engineering, Friedrich-Alexander-University of Erlangen-Nuremberg, Fürther Str. 244f, 90429 Nürnberg, Germany ABSTRACT: The heatpipe reformer provides an allothermal gasification process for the generation of a hydrogen-rich synthesis gas. Heat pipes transport heat from a fluidized bed combustor to a steam-blown fluidized bed gasification reactor. The objective of the Institute of Energy Process Engineering (FAU-EVT) is the generation of hydrogen from the synthesis gas by means of in situ membrane separation in the fluidized bed gasification reactor. This paper presents the current state-of-the-art of the heatpipe reformer (HPR) technology as well as the recent development of the construction of the 100 kW pilot test stand at FAU-EVT and shows the results of a 24 h gasification operation. Lignite was used as feedstock. The results include the synthesis gas and tar composition, the temperature profile of the process, and the sulfur and hydrocarbon concentrations in the synthesis gas. Furthermore, the influence of gasification temperature and pressure on the synthesis gas composition and the influence of the gasification temperature on the tar composition are evaluated. In addition, the anticipated operation points of the HPR, depending on the temperature spread between combustor and reformer temperature and the HPR heat duty, are determined. The generated synthesis gas at gasification temperature of 800−930 °C had a tar content of 2.8−3.6 g/Nm3 and a H2S content of 1000−2000 ppm. Conversion of methane was quite high; its content in the synthesis gas was in the range of 4−7%, resulting in a high hydrogen content of up to 56%.

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mentioned disadvantage of tar contents of 1−10 g m−3. For these applications, the heatpipe reformer is a suitable gasification system as the synthesis gas can be pressurized easier than in other dual fluidized bed concepts, which can be operated only at ambient pressure without pressurizing the combustor. The heatpipe reformer was originally developed at TU Munich during the EU project “BioHPR” (ref. ENK5-CT-200000311).4 The concept was commercialized by Agnion Inc., which built a demonstration plant in Pfaffenhofen, Germany in the scale of 500 kW5 and two commercial plants in the 1 MW scale in Grassau, Germany and Auer, Italy. Figure 1 shows a process scheme of the heatpipe reformer: In contrast to other dual fluidized bed gasification processes, the gasification takes place in a pressurized reformer chamber (1) at 2−10 bar and 800 °C. The pressurized synthesis gas allows usage for SNG synthesis or combustion in engines or gas turbines. The heat for the process is supplied by a combustion chamber (2) located beneath the reformer. Heat pipes (3) transport the heat from combustor to gasifier. To ensure a high heat-transfer coefficient to the heat pipes, both the gasification and combustion take place in fluidized beds. To reduce heat losses and temperature stress on the top flange, the reformer is insulated at the top end (4). Steam (5) and fuel are fed from the top; the fuel enters the fluidized bed with a stand pipe ending in the fluidized bed (6). Table 1 shows typical synthesis gas composition. The heatpipe reformer technology is an ongoing research subject at the Institute of Energy Process Engineering in the course of a subproject of the Bavarian Hydrogen Center

INTRODUCTION Coal gasification is an important process and an ongoing subject of research for higher efficiencies in power generation and the possibility of a CO2-lean usage of fossil fuels. It is also used prior to synthesis steps, as for synthetic natural gas (SNG) production or usage in the chemical industry. Large-scale gasification of coal and lignite uses oxygen-blown fixed bed or entrained flow gasifiers. Concepts for medium- and small-scale biomass gasification propose mainly steam-blown dual fluidized bed (DFB) gasifiers. Entrained-flow gasification operates at high temperatures above 1500 °C because of the usage of pure oxygen as gasification agent. Because of the high temperatures, the synthesis gas is almost tar free.1 In fluidized bed gasification, a bed material (e.g., silica sand, olivine, or fuel ashes) is located in the gasifier. Because of the mixing of the bed material and charcoal, no discrete reaction zones develop. Both pressurized and atmospheric fluidized bed gasification is possible. Allothermal gasification is generally accomplished by dual fluidized bed gasifiers like the dual fluidized bed gasification or the heatpipe reformer. The compact design and lower plant complexity in contrast to entrained flow gasifiers makes fluidized bed gasifiers suitable for biomass or small- to medium-scale coal applications.2,3 Tars pose the most serious problem in gasification: Condensation of tars leads to plugging of piping systems and are a threat to catalysts in synthesis processes because of coking. The highest tar contents occur in updraft gasification followed by downdraft and fluidized bed gasification. The lowest tar contents occur in entrained flow gasification reactors because of the higher process temperature.1 High hydrogen yields make allothermal fluidized bed gasification a promising technology for small- to medium-scale applications and a subsequent usage for chemical syntheses with the above© 2017 American Chemical Society

Received: January 27, 2017 Revised: March 30, 2017 Published: April 3, 2017 4939

DOI: 10.1021/acs.energyfuels.7b00286 Energy Fuels 2017, 31, 4939−4950

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generate a pure sour gas stream. Because of the higher process temperature of this separation process in contrast to the widely used methanol scrubbing, short-chain hydrocarbons remain in the synthesis gas and can be converted on the nickel catalyst. This raises the efficiency of the process chain, which makes it more feasible for small- to medium-scale applications.4,10 In the course of these projects, a 100 kW pilot has been commissioned at FAU-EVT in a 24 h gasification campaign with stable operation. The results from this first gasification campaign with regard to the continuously monitored synthesis gas composition as well as the tar and sulfur content are described in this paper.

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MATERIALS AND METHODS

Gasification of Solid Feedstock with Steam. The stoichiometric reforming of any feedstock with the chemical formula CHmOn with complete conversion to H2 and CO can be described by the following equation:3 CHmOn + (1 − n)H 2O →

⎛m ⎞ ⎜ + 1 − n⎟H 2 + CO ⎝2 ⎠

(1)

Thus, the minimum steam demand, smin, for stoichiometric conversion on gravimetric basis for this reaction can be calculated from the molar masses of feedstock, MCHmOn, and steam, MH2O, as follows:

smin = (1 − n) Figure 1. Scheme of the heatpipe reformer gasification system.3

value

unit

H2 content, dry base CO content, dry base CO2 content, dry base CH4 content, dry base H2O content lower heating value

40−50 15−20 20−25 5−10 20−50 10−15

mol % mol % mol % mol % mol % MJ/kg

MCHmOn

(2)

The gasification is normally operated with excess steam to enhance char and tar conversion; this is defined by the excess steam ratio, σ:

Table 1. Typical Properties of the Synthesis Gas Produced by the Heatpipe Reformer6 property

M H2O

σ=

ṁ s ṁ f smin

(3)

Another important property of the used fuel is the required heat-ofreaction for gasification, Q̇ Δh,r. It is convenient for a practical application to relate this property to the heat input into the gasifier, Q̇ reformer, to yield the specific heat of reaction, q̇Δh,r:3,6

Q̇ Δh ,r Q̇

qΔ̇ h ,r =

reformer

=1−

ṁ sg Hl,sg ṁ rerformer Hl,f

(4)

In this equation, ṁ sg denotes the mass flow of the synthesis gas, Hl,sg the lower heating value of the synthesis gas calculated with the gas composition and higher hydrocarbon content, ṁ reformer the fuel mass flow into the reformer, and Hl,f the known lower heating value of the solid feedstock. Calculation of the Cold Gas Efficiency. The combustor has the most important influence on the cold gas efficiency, ηCG, which is described by the energy content of the produced synthesis gas divided by the input of solid feedstock:

(BHC) and the RFCS Project “CO2freeSNG2.0 - Advanced Substitute Natural Gas from Coal with Internal Sequestration of CO2” (ref. RFCS-CT-2013-0008). Previous publications focused on the main design specifications of the system3 and the performance of the combustor using biomass and lignite as feedstock.7 The goal of the BHC subproject is the in situ separation of hydrogen directly in the reformer pressure chamber using hydrogen-permeable membranes. This allows higher hydrogen yields due to an additional shift of the gasification reactions to the product side as one product is continuously removed.8 Nickel membranes are used for this purpose because of their high stability in these challenging process conditions. The results of the test procedures for these membranes were already described in a previous paper.9 The project “CO2freeSNG2.0” focuses on a process chain for the generation of synthetic natural gas (SNG) by means of catalytic methanation. The Benfield process is used for CO2 and H2S separation between reformer and methane synthesis. The process operates on potassium carbonate solutions used to remove the CO2 and H2S via chemisorption at elevated pressures and temperatures. A second column regenerates the scrubbing agent by depressurizing and stripping with steam to

ηCG =

ṁ sg Hl,sg ṁ f Hl,f

(5)

This mainly depends on the combustor efficiency, ηcomb, which is already discussed in our works on the heatpipe reformer (refs 3 and 7). The following discussion will give only a short overview on the calculations for the cold gas and combustor efficiency. The combustor efficiency, ηcomb, is defined as the heat flux to the reformer by the heat pipes, Q̇ HP, divided by the heat duty of the combustor, Q̇ comb:

ηcomb =

Q̇ HP Q̇

comb

(6)

The combustor efficiency can also be expressed with the following equation using a heat balance around the combustor, where the air inlet temperature, ta, is introduced: 4940


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secondary air is provided by four steel pipes with a nominal diameter of DN25 and is blown into the freeboard radially. The flue gas leaves the combustor at the top with a temperature of about 200−400 °C after preheating primary and secondary air. The temperature monitoring in the reformer is ensured by four thermocouples. The reformer is filled with olivine with a particle size of 0.063−0.180 mm from Sibelco Europe with a bed height of 1.1 m. Pressurized steam and lignite is fed into the gasifier through a stand pipe. This ensures a thorough mixing of steam, lignite, and olivine. Table 2 shows the combustion properties of the lignite used in the experiment. The synthesis gas leaves the gasifier with a temperature of about 500−700 °C through an upper outlet. When heat pipes are operated in a hydrogen-containing atmosphere such as synthesis gas, a small amount of hydrogen permeates through the heat pipe wall and accumulates as noncondensable gas buffer at the upper end. If the hydrogen is not removed during operation, this mechanism causes a rapid deactivation of the heat pipe starting from the condensation zone. Therefore, the heat pipes used in the heatpipe reformer are equipped with a degassing system of hydrogen-permeable nickel membranes, which are placed inside the heat pipes and are flushed with nitrogen. More detailed information on the deactivation problem and the used degassing system can be found in ref 11. Setup of the Process Chain. The heatpipe reformer described above is integrated in a process chain, which includes the fuel, air, and steam supply as well as the conditioning and analysis of the synthesis gas. The piping and instrumentation diagram of this process chain is shown in Figure 4. For the combustion, the fuel input is ensured by a screw conveyor. The inlet of primary and secondary air into the combustion chamber is regulated with mass flow controllers fed by compressed air. The lignite supply to the reformer also takes place by a screw conveyor. To pressurize the implemented fuel, a lock hopper system with three ball valves is installed above the reformer. The fuel is pressurized by means of nitrogen, which also acts as a tracer component for subsequent measurement of the synthesis gas flow. Following the lock hopper system, the steam inlet is mounted. The steam supply is guaranteed by a steam generator and controlled with a Coriolis flow meter. After leaving the gasifier, the synthesis gas is directed into a candle filter, which removes bed material, ash, and char particles at temperatures well above 200 °C. Downstream of the candle filter, the gas and tar analysis of the synthesis gas takes place, as described in the following section. Afterward, a pressure-sustaining valve is mounted, which adjusts the reformer pressure. At the end of the process chain, the flammable components of the synthesis gas are disposed of in the thermal afterburning. Setup of the Gas and Tar Analysis. The flue gas from the combustor is analyzed using a lambda probe Bosch LSM 11, which measures the oxygen content. The CO and NOx content is continuously monitored using an ABB gas analyzer with IR absorption for the components CO, CO2, NO2, and SO2. After the candle filter, an insulated split stream of the synthesis gas is depressurized for analysis of tar content and primary gas components. A sampling port for solid-phase adsorption (SPA) samples is installed in an insulated section to avoid preliminary tar condensation. For each sample, 100 mL of synthesis gas is drawn trough the SPA syringe, which consists of a silica-based amino phase. Later, the samples are eluted and analyzed for 25 individual hydrocarbon compounds with increasing boiling temperatures from benzene to pyrene. The tar analysis is performed with an Agilent 7890A gas chromatography system. After the tar sampling section, the synthesis gas flows trough an impinger bottle to remove heavy tars, water, and other impurities before entering the gas analyzer. An IR analyzer and heat conductivity type ABB AO2020 determines the concentrations of H2, CO2, CO, and CH4. An additional μ-GC 490 system for measuring light hydrocarbons and sulfur components is installed after the ABB gas analyzer. The μ-GC 490 (He as carrier gas) was applied for online measurements of sulfur compounds (H2S, CS2, COS, Ethyl mercaptan, thiophene) and higher hydrocarbons (ethane, ethylene, acetylene, propane, n-butane). A trace-heated (90 °C) stainless steel capillary

(λamin + 1)c p,fgtfg − λaminc p,ata Hl,f

(7)

In this equation, amin denotes the stoichiometric air demand, cp the heat capacity, and Hl, f the lower heating value. Equation 7 is valid only for the ideal case without radiative losses. If a heat balance around the combustor is made, the importance of heat recovery from the flue gas becomes obvious. Without heat recovery, the flue gas would leave the combustor with temperatures of 900−1100 °C, which would result in low efficiencies. Therefore, the heat of the flue gas is used for the preheating of combustion air. The flue gas leaving the preheater will have a temperature of about 400 °C, which is suitable for the generation of process steam for gasification. In Figure 2 the combustor efficiency from eq 7 is plotted for different air−fuel ratios and air preheating temperatures for a

Figure 2. Theoretical combustor and cold gas efficiency. Adapted with permission from ref 3. Copyright 2014 Springer. combustor fluidized bed temperature of 900 °C. To achieve a combustor efficiency of at least 50%, the air has to be preheated to at least 400 °C at an air−fuel ratio of lower than 1.6. With lower air−fuel ratios and high air inlet temperatures, combustor efficiencies of up to 70% are realistic.3,6 If the cold gas efficiency is calculated using these assumptions, a maximum value of 75% can be reached when the char is fully converted and radiative losses are neglected. The most important losses are sensible heat in the remaining flue and synthesis gas. At the Agnion plant in Pfaffenhofen, a maximum cold gas efficiency of 70% was measured at a carbon conversion rate of 80% and a total fuel input of 500 kW.3,5 In previous published results on the combustion chamber, a cold gas efficiency of 75% was predicted for the EVT combustor design.7 This was possible because of the extremely low CO emissions, which allow air ratios as low as 1.2, resulting in high combustor efficiencies above 60% (see Figure 2).

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EXPERIMENTAL SECTION

The following section deals with the setup of the 100 kW heatpipe reformer as well as the process chain, which consists of a sinter metal candle filter and an afterburner for the disposal of the synthesis gas. The gas analysis consisted of an infrared (IR) absorption measurement, a μ-GC for higher hydrocarbons and sulfur components, and SPE samples. The measurement procedures are also introduced after the experimental setup. Setup of the Heatpipe Reformer. Figure 3 shows a sketch of the 100 kW heatpipe reformer at EVT-FAU. The reformer pressure vessel is integrated into the combustion chamber. Eight heat pipes, filled with sodium, transfer the heat, which is needed for gasification, from the combustor into the reformer. The outer lining of the fluidized bed combustor is made of refractory concrete. The primary air inlet is located in the lower part of the combustor. Before entering the combustion chamber, the primary air is preheated. Therefore, the air is lead through an integrated double-walled heat exchanger with a pillowlike structure located at the external wall of the freeboard. The 4941


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Figure 3. CAD sketch of the 100 kW heatpipe reformer at EVT.

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RESULTS The measurement campaign in the presented work had a total duration of 72 h, with 24 h of gasification operation. Start-up of the heatpipe reformer included 24 h of electric preheating to a fluidized bed temperature in the combustor of 250 °C. The combustor was then heated with charcoal to the autoignition temperature of the fuel used for the heating of the process, which lies around 600 °C. The start-up of the combustor to a reactor temperature of 800 °C took another 14 h. After a safe reactor temperature above 800 °C was reached, gasification was started. During the measurement campaigns six operation points were tested including 6 h of pressurized gasification. Table 4 shows the main parameters of the operating points. During operation the fuel input into the gasifier as well as the steam mass flow was kept constant at each operation point, while the heat duty of the combustor was adjusted to maintain steady-state operation at gasification temperatures between 800 and 840 °C. Figure 5 shows the volumetric concentration of the main gasification products, namely H2, CO2, CO, and CH4, as well as the chronological sequence of the different operation points. The concentrations remain stable over the measurement campaign after reaching steady state after approximately 5−10 h, which can can be monitored best with the H2 concentration. The influences of the different process parameters will be discussed in the corresponding sections. The full-load operation point OP 6 is not shown in this figure because it was conducted in a subsequent gasification campaign. Characterization of the Combustor. As discussed above, the combustor has a most critical impact on the efficiency of

Table 2. Properties of the Used lignite (Manufacturer Information) property

value

unit

grain size water content ash content volatile matter fixed carbon carbon content, waf oxygen content, waf hydrogen content, waf nitrogen content, waf sulfur content, waf stoichiometric air demand lower heating value

4−12 19.0 4.0 42 35.5 69.0 24.7 5.0 0.77 0.45 7.1 19.8

mm mass % mass % mass % mass % mass % mass % mass % mass % mass % kg/kg MJ/kg

connects the μ-GC with the outlet of the gas analyzer to avoid adsorption of the synthesis gas components. Intermediate measurements of clean air revealed a fast reduction of sulfur species to less than 3 ppm on the μ-GC columns. Therefore, adsorption effects can be neglected in the range of the detected species within this work. Table 3 summarizes the equipment of the μ-GC together with the calibrated species. In contrast to the applied SPA measurements and the CP Sil 5 CB column, thiophene and benzene (identified by spiking) have been separated by the CP Sil 19 THT column. An overlay of the H2O peak prohibited a measurement of COS. Calibration of the μ-GC was performed with commercial gas samples and was implemented as point-to-point curves. Measurements of several test gases (CS2, COS, ethyl mercaptan, ethylene, n-butane, H2S, O2, N2, CH4, CO, CO2, and thiophene) immediately after disconnecting the μ-GC from the real synthesis gas ruled out a shift of retention times during the campaign. 4942


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Figure 4. Piping and instrumentation diagram of the heatpipe reformer process chain and the gas and tar analysis system.

Table 3. μ-GC Equipment and Parameter Settings column

calibrated species

MolSieve 5A PoraPlot Q

O2, N2, CH4, CO CO2, C2H6, C2H4, C2H2, C3H8, H2S, COS n-butane, CS2, thiophene n-butane, CS2, ethyl mercaptan, thiophene

CP Sil 5 CB CP Sil 19 CB THT

pressure (kPa)

col. temp. (°C)

inj. temp. (°C)

200 250

45 50

100 60

350 200

80 80

100 100

Table 4. Operation Points of the Lignite Measurement Campaign

the process. The most important parameters are the CO emissions with respect to air−fuel ratio and the performance of the integrated air preheater. Figure 6 shows the CO emissions and the air preheating temperatures for a wide range of air−fuel ratios. These values can be used for a calculation of the combustor efficiency by using Figure 2. The maximum combustor efficiency for this system can be calculated as 70− 75% at an air−fuel ratio of 1.2 through Figure 2 as discussed in ref 7. Characterization of the Gasification System. The following section characterizes the reformer in terms of

name

pressure (bar (a))

heat duty reformer (kW)

heat duty combustor (kW)

steam mass flow (kg h−1)

steam ratio (−)

OP1 OP2 OP3 OP4 OP5 OP6

1.3 1.3 1.5 4.0 4.0 4.0

9.9 13.9 17.8 9.9 13.9 29.9

26.2 28.9 34.0 26.8 33.2 50.1

4.0 5.7 7.3 4.0 5.7 12.0

3.0 3.0 3.0 3.0 3.0 3.0

stationarity and temperature profiles. This information is necessary for a prediction of operation points for experiments with higher loads and the performance of the heat pipe degassing system. The gasifier takes some time to reach stationary concentrations. Figure 7 gives exemplary hydrogen concentration and tar content of the synthesis gas. The tar content declines over the course of time while hydrogen concentration is rising. This is probably caused by the buildup of the char content in the 4943


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Figure 7. Behavior of tar and hydrogen concentration after start of gasification.

Figure 5. Chronological sequence of the gas composition during the measurement campaign with the course of the different operation points.

and bed agglomeration. A low temperature difference is also favorable in terms of cold gas efficiency because a lower combustor temperature also raises combustor efficiency.3 Figure 8 shows the temperature behavior over the course of gasification and the above-mentioned temperature difference.

Figure 6. CO emissions and preheating temperatures for different air− fuel ratios and volumetric flow rates through the primary air heat exchanger. Figure 8. Behavior of the reformer temperatures and the temperature spread between combustor and reformer.

olivine fluidized bed.3,6 At the beginning of operation, the synthesis gas is formed by the volatile components of the fuel, while with rising char content and temperature of the reformer, gasification of the coke takes place resulting in higher hydrogen content. As reported in recent literature, e.g., see refs 12 and 13, char from gasification also exhibits catalytic properties, which results in a reduction of the tar content over time. Temperature is a very important parameter of the heatpipe reformer as it has an influence on gasification performance and cold gas efficiency: the temperature spread between combustor and reformer determines the highest gasification temperature because the combustor temperature is limited by ash properties

The heat pipe temperature is calculated by using the internal pressure of the working fluid; bed temperatures are measured with type K thermocouples. The greater temperature difference between combustor bed and heat pipe shows that the higher thermal resistance is located in the combustor. The temperature difference between gasifier and combustor, however, ranges from 60 to 90 K in the experiments, which is an excellent result. At full load operation of 100 kW, a temperature difference in the range of 90−100 K is expected. 4944


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Energy & Fuels This data also proves the performance of the heat pipe degassing system as the reformer shows a nearly isothermal temperature profile. If hydrogen would accumulate in the heat pipes, temperatures decline beginning at the top end of the heat pipes and thus the reformer would be observed within a few hours. Synthesis Gas Composition. Temperature and pressure have an influence on the main synthesis gas components, as shown by the results in Figure 9: at higher pressures, the

Figure 10. Average tar composition of all 25 SPA analyses.

Figure 11. Dependency of the tar content on the gasification temperature.

in Figure 7, only SPA values after reaching stationarity after approximately 7.5 h are displayed. The average tar content rises with decreasing temperature, as expected. GC Analysis of Sulfur and Hydrocarbon Components. The used Agilent 490 μ-GC system is calibrated for some sulfur components, which are H2S, COS, CS2, ethanethiol, and thiophene, as described in the setup section. However, only the results for H2S and thiophene were reliable; the other compounds were either not found (ethanethiol), overlaid by another substance (COS by H2O), or were present in concentrations that were too high (CS2). It is possible that in the complex synthesis gas mixture some substances show a similar retention time, which might explain these analytical problems. The μ-GC also measured higher hydrocarbons from methane to acetylene; the values and separation of the compounds were very reliable. Figure 12 shows the chromatogram of the CP Sil 19 CB THT column for different samples. As discussed above, the CS2 peak increases strongly with runtime. Although the retention time of the peak matched the substance in the calibration gas, the concentration of 0.1−0.2% was found to be too high. Although it is assumed that CS2 is formed from the reaction of H2S with carbon above 800 °C, the concentration stays well below the concentration of H2S because of thermodynamic equilibrium.15−20 It is assumed that mixed substances elute at the same retention time. In contrast to CS2, the enlarged detail

Figure 9. Behavior of the heatpipe reformer synthesis gas composition with changing temperature (at 4.0 bar(a)) and pressure (at 815 °C).

formation of volume-reducing molecules (e.g., methane) is favored over volume-increasing molecules such as hydrogen, which is also visible in the results. These theoretical considerations follow quite accurately the experimental findings shown in Figure 9: methane content rises with the system pressure and declines with rising temperature because of steam reforming. The nitrogen content of the synthesis gas increases with pressure as the lock hopper system needs more nitrogen to pressurize the lignite. Rising temperature causes a higher methane and tar conversion, resulting in higher H2 and CO and lower CH4 and CO2 content. Tar Content and Analysis. Figure 10 shows the average tar composition of the SPA analyses and their standard deviations. Benzene is the component that is present in the highest concentration followed by naphthalene, which is typical for lignite gasification.14 Figure 11 shows the dependency of the tar content on the gasification temperature. The tar content is calculated by the summary of all peaks in the GC analysis that are calibrated. Components, which are not calibrated, are not considered. The tar content could therefore be higher than measured. As shown 4945


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Figure 12. Chromatogram of sulfur components detected with the Agilent 490 μ-GC system and comparison with calibration measurements.

the measured higher hydrocarbons are in accordance with other published data, e.g., from Saw and Pang in ref 21. They found concentrations of 0.1 mol % of ethane and 0.7 mol % of ethylene with pure lignite. The concentrations increased with a higher wood content in the feed. The H2S content remains in the 1000−2000 ppm range during the whole campaign, which is a common order of magnitude for coal gasification.3,5,22 COS measurement could not be performed because the H2O peak overlays the expected COS peak, but there are several publications that indicate a low concentration of COS in lignite-derived synthesis gas under reducing atmosphere.23 Thiophene concentration varied between 10 and 30 ppm, which is rather at the level of biomass gasification than other lignite-derived synthesis gas.23,24 Carbon and Bed Material Balance. The efficiency of the process also depends on a high char conversion in the reformer. Fixed carbon, which presents itself as char, has a much lower reactivity than the volatiles of the fuel and accumulates in the reactor chamber or is pulverized and carried away by the syngas ending up in the hot gas filter. Because of the lower density, an accumulation on top of the olivine bed is also possible. Also, a part of the olivine bed material leaves the reactor with the synthesis gas, accumulates in the filter, and has to be replaced or fed back to the reactor. These parameters are analyzed in the following section by a balance calculation. Figure 14 shows a carbon balance for the operation points LG 1, LG 2, and LG 3. The carbon balance was calculated using the nitrogen tracer and the elementary analysis of the lignite from the manufacturer, which represents an annual average value. It is obvious that the carbon balance is nearly in steady state: the carbon input by the fuel matches the carbon output via the synthesis gas components CO2, CO, and CH4 indicate a nearly complete carbon conversion. The contribution of tars from SPA and higher hydrocarbons from μ-GC analysis to the carbon balance was calculated to be around 1 order of magnitude lower than the contribution of CH4 and thus

at Figure 12 proves that thiophene concentrations revealed minor deviations over the runtime and that separation from benzene was possible. Finally, the baseline for a consecutive air measurement illustrates that adsorption effects in the μ-GC system can be neglected because no thiophene and only a very small amount of CS2 were detected. Nevertheless it should be mentioned that μ-GC analysis was conducted in the outlet gas stream of the gas analyzer, which passed already an impinger bottle and a condenser. Figure 13 shows the concentrations of the compounds measured by the μ-GC system for all operation points in a

Figure 13. Hydrocarbon and sulfur components detected with the “Agilent 490” μ-GC system at operating points shown in Table 4 (chronological sequence).

chronological sequence. The methane concentration was highest with approximately 2−5 vol %, followed by ethylene (0.1−0.8 vol %) and acetylene (0.1 vol %). Ethane and ethylene peaks merged in the ongoing experiments, which made a precise evaluation difficult. In OP 6, the ethane concentration was measured in a second campaign and is more reliable. The content of higher hydrocarbons is generally considered as consequence of the incomplete conversion of products of the pyrolysis step. The concentrations as shown in Figure 13 for 4946


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tracer and the μ-GC nitrogen concentration measurements. Figure 16 shows the resulting plot of the chemical energy

Figure 14. Carbon balance for the operation points LG 1, LG 2, and LG 3 compared with carbon input from lignite. Flow rates were calculated using a nitrogen tracer.

Figure 16. Calculation of the specific heat-of-reaction for the used fuel.

negligible. The value for the carbon content of the fuel is supplied by the manufacturer and represents an annual average value. This might result in uncertainties in the carbon balance. After the experiments, the hot gas filter inventory was weighed and the carbon content of the mixture was quantified, as shown in Figure 15. The filtrate contains bed material, ash,

content of the synthesis gas against the heat duty of the reformer, ṁ reformer Hl, f. According to eq 4, the specific heat-ofreaction can be calculated from the slope of the line of the best fit, which equals 1.20. This yields a specific heat-of-reaction of 20%. This finding is used to calculate the energy balance and the cold gas efficiency for the different operation points, as shown in Figure 17. Beginning from the top, the heat input into the

Figure 15. Hot gas filter inventory analysis and comparison with lignite input. Fuel components are reported according to manufacturer annual average values, after gasification campaign LG 1−LG 5.

and carbon; therefore, the carbon was removed by heating samples to 800 °C for oxidation. The mass loss by this procedure was 23.0%, with a standard deviation of 1.3%, which results in the filtrate masses shown in Figure 15. The comparison of the fixed carbon input with the fixed carbon inventory of the filter makes a more exact calculation of the carbon conversion possible, which results in 85.7%. Gallmetzer et al. also reported a carbon conversion in this order of magnitude in their paper on the 500 kW plant in Pfaffenhofen.5 The relative bed material discharge may be illustrated by comparing the bed material and ash fraction of the filtrate with the integral fuel input. This results in a maximum bed material discharge of 170 g kg−1 Fuel depending on the ash fraction of the filtrate, which could not be assessed. Energy Balance. It is possible to calculate an energy balance from the presented results to evaluate the gasification process in terms of efficiency and potential for optimization. First, the specific heat-of-reaction as introduced in eq 4 is calculated from the measured gas composition and volumetric flows. The volumetric flow rate is calculated from the nitrogen

Figure 17. Energy balance for the different operation points and anticipated energy balance for 1 MW industrial-scale plant; comparison with values of the 500 kW agnion plant in Pfaffenhofen, Germany from Gallmetzer et al.5

reformer (orange) provided by the fuel input is a known value, followed by the heat-of-reaction (red) determined above. The heat for the evaporation of the fuel’s moisture (blue) can be calculated from the known water content in the lignite. The three sensible heat losses (black, stripes) stand for the heat in the synthesis gas, the flue gas, and heat losses. The syngas sensible heat is calculated from the specific heat capacity of the synthesis gas components, the combustor sensible heat from the measured air-fuel-ratio, and preheating temperatures at the respective operation points. The radiative heat loss equals the difference between the sum of the mentioned heat fluxes to the combustor heat duty. It differs between the operation points between 9 and 16 kW. The cold gas efficiency can be calculated by dividing the chemical energy content of the synthesis gas by 4947


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Energy & Fuels the overall heat duty of the process and rises strongly with the heatpipe reformer heat duty. The last bars on the right show a scale-up calculation based on the measured heat fluxes from the described experiments. The sensible heat losses become relatively small in contrast to the heat input to the reformer, resulting in a much higher cold gas efficiency of 70.9%. In contrast to the experiments, an air− fuel ratio of 1.2 and a preheating temperature of 600 °C was considered to calculate the combustor efficiency. With higher preheating temperatures and lower excess steam ratios, a cold gas efficiency of up to 75% is possible. This is consistent with the values of the 500 kW agnion plant in Pfaffenhofen, Germany from Gallmetzer et al., where a cold gas efficiency of 74.3% was reached.5 At the EVT prepilot heatpipe reformer, the radiative heat loss accounts for a high percentage of the overall heat duty of the process, which lowers the cold gas efficiency. This could be overcome by an up-scaling of the process, which would make the sensible heat losses negligible compared to the plant heat duty. Also, the combustor was operated at high air−fuel ratios of 1.5−1.8 for safety reasons. By setting the air−fuel ratio to 1.2−1.4, the sensible heat in the flue gas could be reduced significantly. As shown in Figure 2 and discussed in refs 3 and 7, the cold gas efficiency would then approach the ideal value, which is determined mainly by the combustor efficiency. Mass Balance and Synthesis Gas Yield. The mass balance was calculated using the mass flow from the tracer measurements and the gas analysis of the main synthesis gas components. A comparison with the lignite and steam feed to the reactor results in the steam conversion of 32−34% and a synthesis gas yield of 175%, as shown in Figure 18. The yield reaches values above 100% as steam is also converted into the synthesis gas components.

Figure 19. Yield for different synthesis gas components dependent on the different operation points and for a 1 MW industrial-scale plant.

with the plant scale as cold gas efficiency rises and the fuel input into the combustor becomes lower.

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DISCUSSION The measured synthesis gas properties are in good agreement with literature values, as shown in Table 5. The comparison with other literature data4,5 on the heatpipe reformer reveals similar tar contents and H2/CO ratios. The comparison with other electrically heated or dual fluidized bed gasifiers shows also similar gasification characteristics. H2S content is always in the range of a few 1000 ppm when coal or lignite is used as feedstock. H2/CO ratio is slightly higher with lignite compared to biomass gasification, probably due to a lower content of volatiles. Tar content is in the same order of magnitude as the other literature values. Comparison of the tar contents must always consider the different measuring techniques. Wolfesberger et al. reported in ref 27 that temperatures higher than 850 °C cause a more significant decrease in the tar content, which was not reached in the described experiments. This could lead to a lower tar content. Another measure to further reduce the tar content is an increased char content in the bed inventory.3 The tar yield can give a better comparison for the different gasification technologies. In the heatpipe reformer, the char is converted to a much higher extent than in dual fluidized bed gasifiers, where it is mainly burned in the combustor. This causes an additional dilution of the synthesis gas as steam is converted to CO and H2S, resulting in higher amounts of synthesis gas. As shown in the corresponding section, the heatpipe reformer generates tar yields of 3−10 g kg−1 Fuel. A comparison with recent literature is difficult because few authors give the tar content in terms of a yield. Saw and Pang21 give tar yields for a 100 kW dual fluidized bed gasifier for lignite of 2.3 g kg−1 Fuel, which is lower than in the experiments shown here. These values are in good agreement with a tar yield of 3.2 28 g kg−1 for a 90 kW Fuel calculated from the results of Kern et al. dual fluidized bed gasifier.

Figure 18. Mass balance around the reformer for operation points OP1−OP 3.

This depiction does not consider the fuel mass flow used for process heating. Therefore, Figure 19 gives a comparison of the yield for the respective synthesis gas components. In the first column, the values are based only on the reformer fuel mass flow, resulting in very high yields. When it is based on the complete fuel mass flow, i.e., fuel used for reformer and combustor, it also relies heavily on the cold gas efficiency determined by the process scale. Therefore, the yields are given for two different operation points and the 1 MW industrialscale calculation discussed in the preceding section. Yields rise

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CONCLUSION This paper shows the design and the further development of the heatpipe reformer gasification system. Crucial systems like the combustion management with respect to emission control 4948


Article

Energy & Fuels Table 5. Recent Publications on Synthesis Gas Properties from Allothermal Fluidized Bed Gasifiers

a

gasifier

fuel

heat duty (kW)

H2CO ratio (mol mol−1)

tar content (g Nm−1)

H2S content (ppm)

ref.

HPR HPR HPR El. heated FB El. heated FB DFB DFB DFB DFB DFB

lignite lignite wood pellets lignite wood pellets wood pellets wood pellets lignite lignite hard coal

100 500 500 5 10 100 100 100 100 100

2.8−3.6 2.6−3.0 2.0−2.4 1.4 2 − 1.0 2.4 1.8−2.0 1.7−2.3

3.2−4.8a 2−4a 1.7−2.8a − 0.14−8.72a 4.5−18.5b 9.0b 2.7b 0.8−1b 3.8b

1000−2000 >130 10−23 − − − 104 8000 762 1800

4 5 25 26 27 21 21 28 22

SPA measurements. bMeasurement by gravimetric analysis.

σ = Excess steam ratio

and improvement of the combustor efficiency as well as the heat pipe degassing system were improved and operated. The first results with the feedstock lignite show a moderate tar content in the range of 2.8−3.6 g/Nm3 with SPA measurement. As expected with lignite as feedstock, the synthesis gas contained considerable amounts of sulfur components, with H2S content of 1000−2000 ppm. Ongoing development at FAU-EVT concentrates on the demonstration of the 100 kW SNG process chain with longterm experiments. Scientific development focuses on gas cleaning techniques such as catalytic gas cleaning, hydrogen permeable membranes, and tar and CO2 scrubbing technologies.

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Subscripts

a = Air cg = Cold gas comb = Combustor f = Fuel fg = Flue gas HP = Heat pipe min = Minimum s = Steam sg = Synthesis gas Abbreviations

AUTHOR INFORMATION

Corresponding Author

*E-mail: [email protected]. ORCID

Jonas M. Leimert: 0000-0002-4771-0304 Michael Neubert: 0000-0003-3395-9843

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Notes

The authors declare no competing financial interest.

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REFERENCES

(1) Kaltschmitt, M.; Hartmann, H. In Energie aus Biomasse Grundlagen, Techniken und Verfahren; Kaltschmitt, M., Hartmann, H., Eds.; Springer: Berlin, Heidelberg, 2009. (2) Bridgwater, A. Fuel 1995, 74, 631−653. (3) Karl, J. Biomass Convers. Biorefin. 2014, 4, 1−14. (4) Baumhakl, C.; Karl, J.; Kienberger, T.; Schweiger, A.; Buchholz, D.; Bajohr, S.; Koytsoumpa, E.; Karellas, S. Substitute natural gas from coal with internal sequestration of CO2 (CO2freeSNG); European Commission: Brussels, 2013. (5) Gallmetzer, G.; Ackermann, P.; Schweiger, A.; Kienberger, T.; Gröbl, T.; Walter, H.; Zankl, M.; Kröner, M. Biomass Convers. Biorefin. 2012, 2, 207−215. (6) Metz, T. Allotherme Vergasung von Biomasse in indirekt beheizten Wirbelschichten. VDI-Fortschrittsberichte, Reihe 6: Energietechnik, Nr. 554, 2007; p 134. (7) Leimert, J. M.; Treiber, P.; Karl, J. Fuel Process. Technol. 2016, 141, 68−73. (8) Karellas, S.; Kakaras, E.; Papadopoulos, T.; Schäfer, C.; Karl, J. Fuel Process. Technol. 2008, 89, 582−588. (9) Leimert, J. M.; Karl, J. Int. J. Hydrogen Energy 2016, 41, 9355− 9366. (10) Koytsoumpa, E. I.; Atsonios, K.; Panopoulos, K. D.; Karellas, S.; Kakaras, E.; Karl, J. Appl. Therm. Eng. 2015, 74, 128−135. (11) Leimert, J. M.; Dillig, M.; Karl, J. Int. J. Heat Mass Transfer 2016, 92, 920−928. (12) Klinghoffer, N. Utilization of char from biomass gasification in catalytic applications. Ph.D. Thesis, Columbia University, New York, 2013.

ACKNOWLEDGMENTS This research is carried out in the framework of the Bavarian Hydrogen Center (BHC) joint research program and the EU Project “CO2freeSNG2.0” (ref. RFCS-CT-2013-0008) funded by the research fund for coal and steel (RFCS). The authors acknowledge the support provided by the Bavarian State Ministry of Science, Research and the Arts and the European Union.

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BFB = Bubbling fluidized bed BHC = Bavarian hydrogen center CFB = Circulating fluidized bed DFB = Dual fluidized bed gasifier FB = Fluidized bed HPR = Heatpipe reformer SNG = Synthetic natural gas SPA = Solid phase adsorption

NOMENCLATURE

Latin Letters

a, amin = (Stoichiometric) air demand cp = Specific heat Hl = Lower heating value smin = Stoichiometric steam demand t = Temperature M = Molar mass ṁ = Mass flow Q̇ = Heat flux q̇Δh,R = Specific heat of reaction x = Concentration Greek Letters

η = Efficiency λ = Air−fuel ratio 4949


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Performance of a 100 kW Heatpipe Reformer Operating on Lignite

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