Continuous Hydrothermal Gasification of Glycerol Mixtures

Article pubs.acs.org/IECR

Continuous Hydrothermal Gasification of Glycerol Mixtures: Autothermal Operation, Simultaneous Salt Recovery, and the Effect of K3PO4 on the Catalytic Gasification Martin Schubert,† Johannes B. Müller, and Frédéric Vogel* Paul Scherrer Institut, CH-5232 Villigen PSI, Switzerland S Supporting Information *

ABSTRACT: We have developed a continuous process for the catalytic hydrothermal gasification of wet biomass to synthetic natural gas (SNG). Salts contained in the biomass and released during the liquefaction step are continuously withdrawn in the supercritical salt separation step upstream of the catalytic reactor. The catalytic reactor is operated at temperatures of 400− 450 °C and pressures of 25−35 MPa. In this article we provide a detailed description of the process and demonstrate the proof of concept as well as the process operation characteristics, based on a systematic study of the continuous gasification of aqueous solutions of glycerol with and without K3PO4 with simultaneous salt recovery. Glycerol was gasified efficiently to a methane-rich gas without the formation of tars or char. The gas composition corresponded to the thermodynamic equilibrium. The process could be operated in an autothermal mode, although the large surface-to-volume ratio and the imperfect insulation of the laboratory-scale reactor were responsible for appreciable heat losses along the catalytic fixed-bed. The presence of potassium phosphate, not completely removed upstream of the reactor, led to a shift in the gas composition toward C2−C4 hydrocarbons. However, this effect on the catalyst was reversible.

1.2. Description of PSI’s Catalytic Hydrothermal Gasification Process. At the Paul Scherrer Institute, a continuous catalytic hydrothermal gasification process for the production of SNG has been under development since 2002.18,19 A simplified process flow scheme is depicted in Figure 1. The main steps of this process are •preheating and liquefaction at 350−380 °C: breakup of cells, decomposition of large biopolymers and other macromolecules to smaller organic molecules,4,6,20 release of salts and organically bound heteroatoms such as N, P, and S as inorganic compounds;21 •superheating and salt separation at around 420−450 °C: continuous precipitation and recovery of salts and other minerals before the catalytic reactor22−27 due to the low solubility of salts in supercritical water; •catalytic gasification and methanation at 400−450 °C: final conversion of the organic molecules to mainly CH4 and CO2.20,28,29 The process is operated at pressures of 25−35 MPa ensuring that the water does not evaporate during heat-up. Since salt management was identified as a critical issue for the technical success of sub- and supercritical water biomass gasification technologies,6,7,30 the salt separation step before the catalytic reactor is of particular importance since some salts may act as catalyst poisons such as sulfates28,31−33 or chlorides34,35 and would therefore lead to a deactivation of the catalyst. Thus, the salt separation is necessary to achieve useful catalyst lifetimes. A second aspect is the recovery of the

1. INTRODUCTION 1.1. Hydrothermal Gasification Technology. Hydrothermal gasification can be defined as the conversion of organic substances into a fuel gas in high temperature pressurized water, where the pressure is sufficiently high to keep the water in a liquid or supercritical state. The hydrothermal gasification is typically carried out at temperatures near or above the critical point of pure water (pc =22.1 MPa, Tc = 374 °C). Due to the drastic change in the solvent properties1,2 near-critical and supercritical water provide a unique environment to carry out chemical reactions such as hydrolysis, dehydration, or aldol condensation and aldol splitting.1,3−5 Depending on the temperature, hydrothermal gasification processes may be classified as high temperature supercritical water gasification (typical temperatures above 600 °C) and hydrothermal gasification at moderate temperatures (350−450 °C) in nearcritical and supercritical water.6 The former aims at the production of hydrogen and is typically carried out without catalysts or with the addition of alkali salts as homogeneous catalysts.6,7 Activated carbon is being discussed as an efficient heterogeneous catalyst in the high temperature supercritical water gasification.8,9 This process has been developed up to pilot scale in Germany10−13 and in Japan.8 Hydrothermal gasification at moderate temperatures, on the other hand, aims mainly at the production of methane (or substitute natural gas, SNG). Due to the lower temperatures, (heterogeneous) catalysts are required for the gasification to achieve full conversion and to reach chemical equilibrium, i.e., to obtain a gas mainly composed of CH4 and CO2 with only small amounts of H2.6,14,15 The (catalytic) hydrothermal gasification at moderate temperatures has been developed up to pilot scale at the Pacific Northwest National Laboratory in the U.S.A.16,17 © 2014 American Chemical Society

Received: Revised: Accepted: Published: 8404

February April 21, April 21, April 21,

10, 2014 2014 2014 2014

dx.doi.org/10.1021/ie5005459 | Ind. Eng. Chem. Res. 2014, 53, 8404−8415

Industrial & Engineering Chemistry Research

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Figure 1. Simplified flow scheme of PSI’s continuous catalytic hydrothermal gasification process. PSA: pressure swing adsorption; SNG: synthetic natural gas.

salts as a concentrated brine,23−25 which might be recycled as valuable nutrients and fertilizers after further conditioning. 1.3. Reaction Pathway of the Catalytic Gasification and Methanation. From a macroscopic point of view the hydrothermal gasification of biomass may proceed via an endothermic reforming and a subsequent exothermic methanation step. The maximum theoretical methane yield is obtained when the biomass is fully converted to methane and carbon dioxide (i.e., when neither hydrogen nor carbon monoxide are formed). In the case of the catalytic hydrothermal gasification of glycerol this results in the following exothermic overall reaction stoichiometry:

the investigated metals ruthenium was found to be the most active toward gasification and methane selectivity.40,41 Moreover, supported ruthenium catalysts show a good long-term stability in the hydrothermal environment.28,38 Recent findings obtained by X-ray absorption spectroscopy point to the fact that the gasification is catalyzed by reduced ruthenium (Ru0).34,35,42−44 In particular, operando X-ray absorption spectroscopy studies carried out by Rabe et al.43 and Dreher et al.44 examined the oxidation state of ruthenium during the gasification of ethanol in sub and supercritical water in a continuous flow through setup using a sapphire43 and an aluminum nitride44 capillary as X-ray transparent operando reactor, respectively. They showed that the fresh 2 wt % Ru/C catalyst was in an oxidized state, i.e. the ruthenium was present as Ru+IV species. During heating up in a stream of 5 wt %43 and 7.5 wt %44 EtOH-water mixture the catalyst was reduced to Ru0 at about 150 °C and stayed reduced up to supercritical temperatures of 400 °C,44 the maximum temperature investigated (the continuous experiments were carried out at a constant pressure of 24.544 and 2543 MPa). A recent study by Peterson et al.45 provides further evidence that the pure metallic form of ruthenium is the catalytic active species. According to their study molecular adsorbates are broken down to atomic adsorbates on the ruthenium surface. The adsorbed atomic species are then reformed to CO, CO2, H2, and CH4 according to the thermodynamic equilibrium by rapid scrambling on the catalyst surface. 1.4. Purpose of This Article. Crude glycerol is the principal byproduct of biodiesel production. For example, about 100 kg of crude glycerol are produced per ton of biodiesel.46,47 The amount of produced crude glycerol has increased rapidly due to the rapid increase of biodiesel production. Since 2004, the amount of glycerol produced exceeds the actual consumption.48 Such crude glycerol possesses a very low value due to the impurities contained in it.49 Hence, finding alternative uses of crude glycerol is presently the subject of many research activities.12,36,37,47,49−53 We have performed a systematic study on the continuous catalytic hydrothermal gasification of aqueous solutions of 20

C3H8O3(l) → 1.75CH4(g) + 1.25CO2 (g) + 0.5H 2O(g) ° = −73.7 kJ/mol Δr H298

(1)

According to this stoichiometry a dry product gas composition of about 58 vol % methane and 42 vol % CO2 would be expected and, hence, a molar C:H:O ratio in the dry product gas of 1:2.333:0.833. However, for thermodynamic reasons a small amount of hydrogen remains in the product gas. In fact, an equilibrium calculation (using Aspen Plus) for the gasification of a 20 wt % glycerol solution at 400 °C and 30 MPa yields 55.7 vol % CH4, 41.1 vol % CO2, 3.2 vol % H2, and 0.03 vol % CO, resulting in a molar C:H:O ratio of 1:2.365:0.849 in the dry product gas and a water production of 0.452 mol per mol of glycerol. Note that according to the stoichiometry given in eq 1 water is formed during the (hydrothermal) gasification of glycerol. This is due to the particular C:H:O ratio of glycerol. During the gasification of other feedstocks with differing C:H:O ratios, for example wood, water will be consumed.20 Depending on the process conditions, water may also be consumed in the gasification of glycerol. This is the case for the high temperature gasification in the presence of homogeneous alkali metal or heterogeneous metal catalysts when hydrogen is the desired product gas and hence the water gas shift reaction is promoted.36,37 Supported noble and non-noble metal catalysts have been investigated for the hydrothermal gasification.14,15,38−40 Among 8405



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Figure 2. Process flow schematic of PSI’s continuously operated laboratory plant for the catalytic hydrothermal gasification of biomass. Legend: E, electrical conductivity meter; F, gas flow meter; FC, flow controller; H, electrical heater; P, pressure transmitter; PC, pressure control valve; PI, pressure gauge; T, thermocouple; TC, temperature controller.

monitor the feed flow rate. Using a Novados N−P31 singlepiston high pressure pump (Bran & Luebbe, max. throughput 1 kg/h) the feed is pressurized to 25−30 MPa and pumped through the electrical preheater (stainless steel 1.4435, length 1.70 m, 12 mm ID and 18 mm OD) where it is heated up to temperatures of 300−370 °C. After passing a heated transfer tube (stainless steel 1.4435, length 0.335 m, 2.4 mm ID, and 6.35 mm OD) the feed enters the salt separator where it is superheated to fluid temperatures up to about 450 °C very rapidly by mixing with the hot fluid inside the heated salt separator. The salt separator is designed to be the hottest part of the test rig. A detailed description of the salt separator was published elsewhere.24 A small stream is withdrawn at the vessel’s bottom as a concentrated brine and cooled in a tube-intube cooler. The flow rate of this stream is controlled within a range of 0−3.2 g/min using a mass flow controller (Bronkhorst LiquiFlow). The main part of the feed stream leaves the separator vessel at the top exit port through a transfer tube (optionally heated, stainless steel 1.4435, length 0.145 m, 2.4 mm ID, and 6.35 mm OD) to the reactor. The reactor tube (stainless steel 1.4435, length 1.40 m, 12 mm ID, and 18 mm OD) is filled with a fixed-bed of a 2 wt % Ru/C catalyst (BASF). The catalyst, granules of about 1−2 mm size with an irregular shape, is used as received, without prior reduction or

wt % glycerol and 20 wt % glycerol with 0.05 or 0.1 mol/kg K3PO4. Potassium phosphate as an additive was chosen because K3PO4 was considered a relatively inert salt with respect to side reactions and corrosion, and because it represents the phosphorus fraction of biomass. In this article, we describe the realization of the proposed process at laboratory scale and discuss the effect of salts on the gasification of glycerol solutions containing K3PO4.

2. EXPERIMENTAL SECTION 2.1. Description of the Continuously Operated Laboratory Plant. All experiments described herein were carried out in a continuously operated laboratory plant first described by Waldner.29 Since then it has undergone several improvements, described in the following. The continuously operated laboratory plant consists of seven sections: Feeding section, preheating, superheating and salt separation, catalytic reactor, cooling, pressure control and letdown, and gas−liquid phase separation. A simplified flow sheet of the plant is depicted in Figure 2. The feed section consists of two containers, one for the feed solution to be studied (glycerol, glycerol + K3PO4, crude glycerol), the other for deionized water for flushing the plant after an experiment. The containers are placed on a balance to 8406



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analysis. The flow rate of the gaseous brine effluent was measured manually using a bubble flow meter. A certain liquid level in the phase separator of the reactor effluent was necessary to ensure a continuous operation. Due to fluctuations in the flow rate and the gas to liquid ratio of the reactor effluent during gasification, it was difficult to keep a defined liquid level in the phase separator. However, to manually control the liquid level a transparent riser tube (standard silicon lab hose) was connected as a siphon to the bottom outlet of the phase separator. In contrast to pure salt separation experiments carried out in an earlier study using the same setup,24,25 in most cases of the gasification experiments, in particular for the gasification with simultaneous salt separation, no real steady state concerning pressure, the effluent conductivities and the effluent mass flow rates could be reached. Thus, to minimize errors in the material balances, in most runs average samples of the liquid effluents were taken in polyethylene flasks during at least 30 min. If not stated otherwise, all mass balances were calculated from the average samples. In addition to these sampling cycles, also small momentary samples were taken. If not stated otherwise the gas sampling and the measurement of the gas flow rate at the brine effluent were carried out during taking the average samples. 2.3. Sample Analysis. Online and off-line gas analysis was performed by gas chromatography with an Agilent HP6890 gas chromatograph. Helium was used as carrier gas. A two column switching system (a HP-Plot Q and thereafter a HP-Plot Molecular Sieve 5A column), connected to a thermal conductivity detector (TCD), allowed the separation and detection of CO2, H2, CH4, and CO (as well as Ar/O2 and N2). Other hydrocarbons were separated using an HP-1 hydrocarbon column and detected using a flame ionization detector (FID). The GCs were calibrated using four different calibration gases (see the Supporting Information). Before online or offline gas analysis, a calibration gas (mostly mixture 1) was injected several times to control the stability of the calibration. Gas chromatography coupled to a mass spectrometer (GCMS) was used for a qualitative identification of compounds in liquid product phases. The GC was also an HP6890 using Helium as carrier gas and an M+N 726420.60 Optima delta 3 column. Compounds were identified by comparison of the measured mass spectra with mass spectra from a database (Masslib, Version A.03.00, Hewlett-Packard). Inductively coupled plasma−optical emission spectrometry (ICP-OES) was used for quantitatively analyzing the aqueous samples for the elements K, P, S, Fe, Ni, Cr, Ti, Al, V, and Ru. The spectrometer was a Cirros ICP-OES device (Spectro). This device was calibrated with different dilutions of multi element standard solutions (purchased either from Kraft or from Merck). Ion chromatography was used to measure the concentration of PO43− in the aqueous effluents. The chromatograph was an HPLC Summit system (Dionex) equipped with a Metrosep A Supp 5 analytical anion separation column (Metrohm) and an ASRS 300 anion suppressor (Dionex). Detection was performed with an ED 50 electrochemical detector in the conductivity mode. The Total Carbon (TC), the Total Organic Carbon (TOC) and the Total Inorganic Carbon (TIC) of the liquid effluents were measured using a vario TOC cube (Elementar). For determining the TC, the samples were directly injected into a quartz glass reactor where the samples were burnt in a constant oxygen carrier gas flow (200 mL/min, gas purity 4.3) at 850 °C.

grinding. For gasification experiments the reactor is heated to about 400 °C. Having passed a cooler (stainless steel 1.4435, coiled tube, length 3.8 m, 2.4 mm ID, and 6.35 mm OD, manufactured in our workshop) after the reactor, the fluid is depressurized to atmospheric pressure in two steps using a high pressure control valve (Kämmer) and a pressure controller (Flowserve), followed by a spring-loaded relief valve (Sitec, Switzerland). In a phase separator made of borosilicate glass, gaseous products are then separated from the remaining aqueous phase. The electrical conductivities of the phase separator effluent and of the brine are monitored online with a conductivity meter (WTW Cond 340i) using a flow-through cell fabricated in-house. The brine and the phase separator effluent are collected and weighed on laboratory scales. All mass flow rates are calculated from the scale signals and are reported as mean value and its standard deviation. All temperatures, pressures, scale signals, electrical conductivities of the effluents and the amount of gas produced as well as the gas composition are monitored online. The laboratory plant can be operated by remote control using a LabView-based control program. 2.2. General Procedure and Parameters for the Gasification Experiments. For all experiments, a pressure of 30 ± 0.5 MPa was maintained. Each experiment was started by pumping deionized water and heating up until the desired temperatures in the preheater, the salt separator and the reactor of the respective experiment had been reached. After heating up with deionized water, samples of the salt separator effluent and the reactor effluent were taken to have a system blank value, and the feed was then switched to the respective glycerol− water or glycerol-K3PO4-water mixture. The main experimental parameters are listed in the following: • glycerol concentration: ∼20 wt % (TOC = 75−81 g/kg) • K3PO4 concentration: 0, 0.05, or 0.1 mol/kg (corresponding to approximately 0, 1, and 2 wt %) • pH of the feed solution: 7.0 (without K3PO4) or 12.5−13 (with K3PO4) • mass flow rate of the feed: 16.4−18.6 g/min • mass flow rate of the brine effluent: 1.0−3.0 g/min • mass flow rate of the reactor effluent: 11−14 g/min • temperature set point of the preheater: 300−370 °C (standard experiment 300 °C) • temperature set point of the salt separator: 430−500 °C (standard experiment 470 °C) • temperature set point of the reactor: off, or 400 °C (standard experiment 400 °C) • Weight hourly space velocity (WHSV): 2.3−2.7 gorg/ (gcat·h) The detailed experimental parameters are given in Table S1 of the Supporting Information. The chemicals used in this study are listed in the Supporting Information. The water used for flushing the plant, for preparing the glycerol and glycerol-salt solutions, and for diluting the crude glycerol was deionized water (electrical conductivity ≤5 μS/cm) taken from the in-house grid. As also the brine effluent contained some gases, they were separated after depressurization in a small phase separator made of borosilicate glass. The liquid part of the brine effluent (further referred to as brine or brine effluent) was continuously collected on a scale. The gaseous part of the brine effluent (further referred to as gaseous brine effluent) was collected in gastight sampling bags (purchased from SKC) for off-line GC 8407



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The evolving CO2 was quantified with a nondispersive infrared (NDIR) detector. For determining the TIC, the (aqueous) samples were injected into a 10 wt % solution of H3PO4 and the evolving CO2 was stripped with the oxygen carrier gas flow. The TOC of the aqueous samples was obtained as the difference between the TC and the TIC. All liquid samples were diluted with deionized water to match the calibrated range of each analytical method. More detailed information regarding the analytical methods can be found in the Supporting Information. 2.4. Data Evaluation. For all experiments the individual mass flow rates of organic and inorganic carbon, potassium and phosphorus (or phosphate) in the liquid effluents were calculated with eq 2:

ṁ i = ciMim/ ̇ ρ

Y ′CH4 [gCH − C /gC] = 4

Cout /C in[%] =

where ṁ i is the mass flow rate of carbon, potassium, or phosphorus, ci is the molar concentration of each element or ion, Mi is the molar mass of the respective element or ion, and ṁ and ρ denote the total mass flow rate and the density of the influent or effluent stream, respectively. The density was measured for representative feed and effluent solutions obtained during the gasification runs by weighing portions of 10 mL of the respective solution. The densities of those representative solutions are given in the Supporting Information. The total volumetric flow rate of the gaseous reactor effluent was determined from the plot of the cumulative gas volume produced during the respective experiment vs time. The total molar gas flow rate at standard pressure and temperature (0 °C, 0.1013 MPa) was calculated using the ideal gas law. Hence the mass flow rate of each individual gaseous component which was detected was calculated using eq 3: (3)

where ṁ j is the mass flow rate of the individual gas component, ṅGas is the total molar gas flow rate, yi is the mole fraction of the individual gas component (which corresponds to its volume fraction determined online using gas chromatography), and Mj is the molar mass of the individual gas component. The total mass flow rate of the gaseous reactor effluent is thus a sum of the mass flow rates of the individual gas components. The total carbon mass flow rate in the gaseous effluent is thus: ṁ C,Gas =

̇ yj zjMC ∑ nGas

(4)

j

GECReact[%] = CGas/C Feed × 100% React ṁ C,Gas Brine Brine ṁ CFeed − (ṁ C,Gas ) + ṁ C,Liquid

(6)

× 100% (7)

3. RESULTS AND DISCUSSION 3.1. General Results: Total Mass and Elemental Balances. For a typical experiment the evolution of the product gas composition in the reactor effluent and the evolution of the electrical conductivities of the effluents, which are a measure of the total salt concentration, are depicted in Figure 3. A steady-state concerning the product gas composition and the effluent conductivities was reached in both runs of this experiment. The product gas composition of the gasification without salt consisted of 56.6 ± 0.4 vol % CH4, 41.8 ± 0.2 vol % CO2, and 1.2 ± 0.4 vol % H2 and was thus close to the chemical equilibrium. It should be noted that the equilibrium gas composition will be different from the one expected for glycerol (according to eq 1), if the C:H:O ratio of

where ṁ C,Gas is the total carbon mass flow rate in the gaseous effluent, zj is the number of carbon atoms of the individual gas component, and MC is the molar mass of carbon. The carbon gasification efficiency (GEC), the methane yield (Y′CH4), and the carbon balance (Cout/Cin) are important parameters for characterizing the gasification. These parameters were determined as follows:

=

ṁ CFeed

Brine Brine ṁ CFeed − (ṁ C,Gas ) + ṁ C,Liquid

where GEReact is the carbon gasification efficiency in the reactor, C CGas/CFeed is the ratio of carbon in the product gas to the carbon in the reactor feed; ṁ React C,Gas is the mass flow rate of carbon in the product gas; ṁ React is the mass flow rate of carbon in the C Brine feed; ṁ Brine C,Gas and ṁ C,Liquid is the mass flow rate of carbon in the React gaseous and liquid brine effluent, respectively; ṅCH is the 4,Gas molar flow rate of methane in the product gas; MC is the molar mass of carbon; Cout/Cin is the mass flow rate of all effluent carbon to the mass flow rate of the influent carbon (carbon balance); ṁ eC is the mass flow rate of carbon in the individual effluent where e denotes the respective effluent. If the chemical equilibrium is reached in the gasification reaction (compare gas composition at thermodynamic equilibrium presented in section 1.3), a feed stream of 1 kg/ h of a 20 wt % glycerol solution would result in 60 g/h of methane, corresponding to a methane yield (Y′CH4) of 0.58 gCH4‑C/gC. Note that in the following, if not stated otherwise, GEC denotes the carbon gasification efficiency in the reactor (GEReact C ) and the superscript (React) is omitted. Thus, GEC does not represent the overall gasification efficiency of the process. We used the carbon gasification efficiency also as proximate measure for the activity of the catalyst. The overall gasification efficiency of the process can be found in Table S6 of the Supporting Information of this article. The procedure is analogous for the gaseous salt separator effluent, with the difference that gas analysis is carried out offline and the volumetric gas flow rate had to be measured manually. Equation 7 was analogously applied for calculating the overall recoveries of potassium and phosphorus (and phosphate) by using the mass flow rates of either potassium or phosphorus instead. For calculating the recoveries of carbon, potassium or phosphorus in either the brine or the reactor effluent, the sum in the numerator of eq 7 was replaced by the mass flow rate of the individual element (or ion) in the respective effluent. The assumptions and equations which were applied for the error estimation of the mass balances, the carbon gasification efficiency as well as the methane yield can be found in the Supporting Information.

(2)

ṁ j = nGas ̇ yj Mj

∑e ṁ Ce

nCH · MC ̇ React 4,Gas

× 100% (5) 8408



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whereas 57.5% glycerol conversion and 10.7% coke yield was found under the same conditions without NaOH. Under supercritical conditions (500 °C and 34.5 MPa) Antal and coworkers58 observed 32% conversion of glycerol to acrolein, acetaldehyde and gaseous products (H2, CO, CO2, CH4, C2H4, and C2H6) without a catalyst at a residence time of 1.5 min. These conditions are not too far from the conditions inside our salt separator (30 MPa, peak temperature of 460 °C and a similar mean residence time). However, in this study, we only analyzed the effluents (or the reactor inlet) for dissolved carbon (TOC and TIC). Hence, we can neither comment on liquid phase products in the brine effluent, nor can we estimate the true C: H: O ratios in the reactor feed. Salt separation and the recovery of a concentrated brine was achieved in run G14b (gasification of glycerol with K3PO4). This is indicated by the conductivity of the brine effluent which is higher than the conductivity of the feed solution by a factor of 3−4. After about 4 h on-stream with the glycerol-K3PO4 mixture, salt was detected in the reactor effluent, as indicated by an increase of the reactor effluent’s conductivity. However, the mean conductivity of the reactor effluent was below the conductivity of the feed solution. Hence, the total salt concentration in the reactor effluent was below the salt concentration of the feed solution. The complete mass and elemental balance (for the main elements C, K and P) as well as the material flows of experiment G14 are depicted in Figure 4. As can be seen from Figure 4 the brine effluent contained ca. 10−12% of the feed carbon. Due to the good solubility of organic molecules in supercritical water,6 most of the carbon flow is split in the salt separator according to the mass flow ratio between the brine effluent and the feed solution. Thus, the carbon concentration in the brine effluent was quite high in all gasification experiments. In fact, the total carbon concentration in the liquid brine effluent was 68.4 and 49.8 g/kg during the runs without K3PO4 (run G14a) and with K3PO4 (run G14b), respectively; whereas the total carbon concentration was 76.9 and 78.9 g/kg in the feed solutions of run G14a and run G14b. The carbon mass flow rate was smaller than one would expect from the mass flow ratio “brine effluent: feed solution”, even

Figure 3. Experiment G14: Evolution of the product gas composition and the electrical conductivities of the effluents for the gasification of 20 wt % glycerol (run G14a) and 20 wt % glycerol with 0.05 mol/kg K3PO4 (run G14b). Double arrows indicate the periods of taking an average sample (see also Figure 4).

the organic matter in the reactor feed is different from the C:H:O ratio of the organic matter in the feed stream. This would be the case if the organic matter withdrawn in the brine effluent exhibited a C:H:O ratio different from the one of the feed. This was most likely the case in our experiments, as the glycerol underwent some decomposition already in the preheater and in the salt separator as shown in our companion article54 and by several other studies.9,52,55−57 At 300 °C, 34.5 MPa and 16 s residence time, Ramayya et al.55 observed 6% glycerol conversion with acrolein and acetaldehyde as major products in the presence of acid. Interestingly, in the presence of a base (NaOH), Müller and Vogel57 did hardly observe any conversion of glycerol at 350 °C and 30 min residence time

Figure 4. Experiment G14: Mass flow rates, material balances and distribution of the elements C, K, and P into the different effluent phases at steady state conditions. Italic style and gray color refer to run G14b (gasification of glycerol + 0.05 mol/kg K3PO4). 8409



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Figure 5. Experiments G3 and G4: Influence of the reactor inlet temperature (by applying different salt separator set point temperatures) on the gasification of 20 wt % glycerol without salt. G3-salt separator set point 500 °C. G4a and G4b-salt separator set point 470 °C. G4c-salt separator 450 °C. G4d-salt separator 430 °C. All lines in these diagrams are guides to the eye. (a) Reactor temperature profiles. Full symbols: reactor set point 400 °C. Open symbols: reactor heating switched off. WWa and WWb: Temperature profiles of water at a salt separator set point of 500 °C and a reactor set point of 400 °C (WWa) or at a salt separator set point of 470 °C with the reactor heating switched off (WWb). For both runs with water, the feed and the brine effluent flow rates were 16.5 and 2.5 g/min, respectively. (b) Product gas composition at steady state of each run. G3a and G4a: Product gas composition of the runs with the reactor set point at 400 °C.

gasification efficiencies of 108 ± 7 and 103 ± 7% for run G14a and run G14b, respectively. The higher content in inorganic carbon of run G14b resulted from an increase of the pH value due to the presence of salt in the liquid reactor effluent of run G14b. In contrast to the brine effluent, for all gasification experiments the reactor effluent appeared as a colorless and odorless aqueous liquid. Thus, neither tars nor char were formed in the catalytic gasification. The separation efficiency for phosphorus (and phosphate), i.e., the recovery in the brine effluent, was higher than the separation efficiency for potassium. The opposite was observed for the respective recoveries in the aqueous reactor effluent. This was the case for all gasification experiments, which were carried out at conditions where efficient salt separation could be achieved, i.e., with salt separator set point temperatures of 470 °C. The difference in the recoveries of potassium and phosphorus (or phosphate) may be explained by the protonation (and perhaps condensation) of the PO43− in supercritical water.24,64,65 This protonation might be even more pronounced for the gasification experiments, as organic acids are produced during the hydrothermal degradation of glycerol (compare the pH of the brine and the reactor effluent in run G14a and run G14b). It is known that simple organic acids, e.g., acetic acid, are produced during the hydrothermal degradation of organic molecules.15,66,67 The organic acids are partially neutralized by the alkaline potassium phosphate salt. Generally, these observations for the material flows hold for all experiments with the aqueous solutions of glycerol (with or without K3PO4). Although the molar C:H:O ratio of the reactor feed most likely differed from the molar C:H:O feed ratio of glycerol (compare discussion above) the dry product gas of the reactor effluent showed a mean molar C:H:O ratio of 1:2.343:0.836 for all gasification experiments of glycerol with or without K3PO4. There was no significant difference in the molar C:H:O ratio between the runs with or without K3PO4. The

when accounting for the gaseous brine effluent. However, the splitting of the carbon stream inside the salt separator according to the mass flow ratio “brine effluent: feed solution” is still a good estimate. In the presence of K3PO4, the carbon concentration in the liquid brine effluent decreased whereas a significantly higher carbon mass flow rate was found in the gaseous brine effluent (0.49 vs 0.79 g/h; refer to Figure 4 and Table S11 of the Supporting Information). This effect may be attributed to an enhanced gasification catalyzed by the salt (or more specifically by K+ of the K3PO4) and is discussed in more detail in our companion article.54 Moreover, this finding is consistent with results in the literature.7,59,60 For example, Chakinala et al.61 observed complete glycerol gasification at 600 °C, 25 MPa and 5 s residence time in the presence of 0.5 wt % K2CO3. In contrast, the gasification efficiency was only 66% without K2CO3 under the same conditions. Furthermore, in the presence of K3PO4 only small amounts of CO were present in the gaseous brine effluent, the gas composition shifted toward CO2 and H2 which is attributed to a catalytic effect of K+ on the water gas shift reaction.62,63 It should be noted here that the brine effluent showed a yellow color and a second dark yellowbrownish oily liquid phase after some time during the gasification of glycerol without K3PO4. This might be an indication for tar formation. However, in none of the experiments any evidence for char or coke formation such as plugging or a lack of carbon was observed. On the other hand, when gasifying glycerol in the presence of K3PO4 the effluent appeared only light yellow and no second oily phase was formed. Thus, in the presence of K3PO4 no tars were formed. All carbon detected in the aqueous reactor effluent of run G14a was inorganic carbon, i.e., dissolved CO2 and its ions (see also Table S8 of the Supporting Information). Likewise, about 66% of the carbon detected in the aqueous reactor effluent of run G14b was inorganic carbon. Thus, the gasification can be considered complete for both runs, which is also reflected by 8410



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There was no difference in the product gas composition of these runs. Thus, the product gas composition of these runs is summarized as “G3a and G4a” in Figure 5b. Especially for the runs G4c and G4d, the product gas composition shifted away from the equilibrium composition having less methane and a considerable amount of C2 to C4 hydrocarbons. The appearance of the higher hydrocarbons may thus be attributed to the lower reactor temperatures compared to the runs with the heated reactor or to a higher salt separator set point temperature. Due to lower reaction rates, the residence time in the catalytic bed is not long enough for the organic matter to be fully converted to C1 components at lower reactor temperatures. Apparently, the reactions responsible for the decomposition and reforming (“gasification”) are more affected by lower reactor temperatures than the water gas shift reaction and the methanation, as no CO was detected in the product gas. This can be explained if one assumes that the steam reforming exhibits a higher activation energy in hot pressurized water in the presence of a ruthenium catalyst than the methanation and the water gas shift reaction.15 Another trend can be seen for the hydrogen. For lower reactor temperatures, less hydrogen was produced and the amount was below the detection limit for the runs G4b-G4d. This trend was also observed for the experiments G12a and G12b (see Figure 6a). As the hydrogen production is thermodynamically favored at higher temperatures,7,29,69 lower temperatures lead to less hydrogen in the product gas. Since a product gas of equilibrium composition was obtained for salt separator set points of 470 and 500 °C without actively heating the reactor for the gasification of 20 wt % glycerol (compare Figure 5b), it was interesting to see if this was also the case when gasifying a solution of glycerol and salt. Thus, in experiment G12, solutions of 20 wt % glycerol (G12a and G12b) and 20 wt % glycerol with 0.05 mol/kg K3PO4 (G12c and G12d) were gasified at a salt separator set point temperature of 470 °C, switching from an actively heated reactor (set point 400 °C, runs G12a,c) to a nonactively heated reactor (reactor heating switched off, runs G12b,d) for both feed solutions. The evolution of the product gas composition in the reactor effluent and the evolution of the electrical conductivities of the liquid effluent are depicted in Figure 6a. Compared to the experimental run G4b (Figure 5b), 0.8 ± 0.1 vol % H2 was found in the product gas when the reactor was not actively heated (run G12b). However, the amount of hydrogen in run G12a (reactor heated to 400 °C) was significantly higher, i.e., 1.5 ± 0.3 vol %. Thus, the experimental runs G4a-b and G12a-b show the same trend. The differences of the runs G4b and G12b are attributed to an optimized GC analysis method allowing a more reliable detection of hydrogen concentrations below 1 vol % H2. In order to compare the gasification of 20 wt % glycerol and 20 wt % glycerol with of K3PO4, the temperature profiles of the reactor of experiment G12 are shown in Figure 6b. Clearly, the endothermic and exothermic reaction regions are seen in all runs of experiment G12. For the runs with the heated reactor (G12a and G12c), the temperature profiles rejoin the temperature profile of water under the same heating and flow conditions. Interestingly, in the presence of salt (runs G12c-d), the endothermic part of the temperature profile is more pronounced. This is also the case for the exothermic part as the temperature of run G12d, compared to run G12b, increased to at least the same value in the exothermic part, and remained at slightly higher temperatures until almost the end of the catalytic

observed molar C:H:O ratio was close to the one expected from eq 1 or, more precisely, close to the one expected from the thermodynamic considerations (refer to section 1.3). This is a strong indication that water was indeed produced during the gasification as expected from eq 1. Complete mass and elemental balances were determined for each experiment. These values are given as tabular data in Table S5 of the Supporting Information. In most cases, the total mass balance and the carbon balance closed well within the experimental error. In some cases the carbon balance did not close whereas the total mass balance closed well to 100%. This might be attributed to small but undetected leaks in the product gas line, e.g., in the phase separator. A nonclosing overall salt recovery can be attributed to not having reached a real steady state in that respective experimental run even when taking long average samples. More detailed results regarding the material flows, i.e., the overall carbon gasification efficiency of the process, the product gas compositions of the brine and the reactor effluent, the molar flow rates of the individual gases in the respective gaseous effluent, the pH values of the liquid effluents, the carbon mass flow rates in the liquid and gaseous effluents as well as the total mass flow rate of the gaseous effluents, are given as tabular data in the Tables S5−S11 of the Supporting Information. 3.2. Autothermal Operation and Influence of the Reactor Inlet Temperature on the Gasification. Ideally, the hot reaction mixture entering the reactor after the salt separator would provide enough heat to drive the reaction without the reactor being heated, because the overall reaction (1) is slightly exothermic. Thus, a set of experiments was performed by varying the salt separator temperature set point to estimate the minimum reactor inlet temperature for which still a product gas near the chemical equilibrium composition was reached, see Figure 5. There are two reaction zones, which are clearly visible in the temperature profiles of runs G3b and G4b (Figure 5a). From a macroscopic point of view, the endothermic part may be attributed to a reforming step29 of the glycerol, while the exothermic part may be attributed to the methanation reactions. However, the interpretation of these temperature profiles is complicated by the fact that the enthalpies of the different reactions are superimposed by the negative mixing enthalpy of the supercritical reaction mixture with the evolving gas.68 These endothermic and exothermic zones within the reactor were less pronounced for a lower reactor inlet temperature (see temperature profiles of the runs G4c and G4d). This may be explained as follows: A lower reactor inlet temperature decreases the rates of all reactions. Particularly, the rate of the endothermic reforming reaction might be strongly affected by a lower reactor fluid temperature leading to a less pronounced endothermic feature in the temperature profile. Thus, the reforming reactions do not proceed at once when the organic matter reaches the active zone of the catalyst bed. All reactions may be rather dispersed along the catalytic bed. Hence, the different reactions are superimposed, and so are the heats of reaction, leading to less pronounced temperature features in the profile. Finally, decreased reaction rates also lead to slower gas evolution which in turn leads to less heat absorption by the mixing of the gas and the supercritical fluid. For the runs with the heated reactor (set point 400 °C), the product gas composition was close to the chemical equilibrium. 8411



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observed at the beginning of run G12d with the reactor not actively heated (Figure 6). However, the salt breakthrough was not caused by the lower average reactor temperature but by the saturation of the catalytic bed with salt, as a salt breakthrough was observed for all gasification experiments in the presence of K3PO4 (compare also Figure 3). The average conductivity of the reactor effluent was lower than the conductivity of the feed solution, whereas the brine effluent had a 3−4 fold higher conductivity than the feed. Up to 69% of potassium and up to 86% of phosphorus were recovered in the brine. The remaining fraction in the reactor effluent was 24% for potassium and 4% for phosphorus (see Table S5 of the Supporting Information). In the presence of salt in the reactor (see for example runs G12d and G14b in Figure 3 and Figure 6), the gas composition shifted to higher hydrocarbons, i.e., C2−C4 hydrocarbons, summing up to a total amount of about 1 vol %. Generally, the occurrence of the higher hydrocarbons was accompanied by a slight decrease in the methane concentration. However, this was not the case for experiment G12 in which, in the presence of K3PO4, the concentration of CH4 and CO2 remained constant, whereas no hydrogen was detected in run G12d. This behavior may be attributed to the lower temperatures in the reactor (the reactor was not actively heated during the run G12d) and to the presence of salt in the reactor. However, the occurrence of higher hydrocarbons is clearly attributed to the remaining salt in the reactor feed (i.e., salt which was not removed in the salt separator). Apparently, the higher hydrocarbons occurred after a certain amount of the catalyst bed had been saturated with remaining salt. This might be explained as follows: At the saturation level where the higher hydrocarbons were detected the remaining amount of “free” catalyst was not high enough, i.e., the space velocity on the “free” catalyst was too high to fully convert these hydrocarbons into CH4 and CO2. We investigated the effect of the space velocity for the gasification of glycerol in a different experimental setup.70 Full gasification of pure glycerol and a gas of equilibrium composition was observed up to a weight hourly space velocity of 11 gorg/(gcat·h). A WHSV of 18.7 gorg/ (gcat·h) led to a decrease of the gasification efficiency to about 80% and to an overall fraction of C2−C4 hydrocarbons of around 4% in the gas phase (methane dropped to 40%, hydrogen increased to 12%, no CO was detected). Switching back to a WHSV of 3.7 gorg/(gcat·h) showed the full reversibility of this effect.70 Note, however, that the catalyst was not permanently deactivated by the salt. The catalytic performance could be fully recovered by flushing the catalyst bed with water after the experiments. Gasification runs without salt (e.g., G12a,b, G13a, G14a and so forth) showed no higher hydrocarbons in the product gas. A permanent catalyst deactivation was, however, observed during the gasification of crude glycerol which served as a real substrate.22,71 Since no ruthenium was detected (using ICP-OES analysis) in the reactor effluents of the gasification experiments, specifically in the gasification of crude glycerol, massive leaching of the catalyst was ruled out as cause for deactivation of the catalyst during the gasification of crude glycerol.22,71 Thorough catalyst characterization after gasifying crude glycerol led to the assumption that the deactivation was most likely caused by fouling with some carbonaceous species.71 The effect of alkali metals on ruthenium catalysts was studied for the water gas shift reaction, the hydrogenation of CO and the ammonia synthesis.72−77 Panagiotopoulou and Kondarides74 found that doping TiO2 supported Pt, Ru, and Pd

Figure 6. Experiment G12: (a) Evolution of the product gas composition and of the electrical conductivities of the effluents during the gasification of 20 wt % glycerol (G12a and G12b) and 20 wt % glycerol with 0.05 mol/kg K3PO4 (G12c and G12d). For both feed solutions, the reactor was either heated to 400 °C (G12a,c), or not actively heated (reactor heating switched off, G12b,d). (b) Temperature profiles of the reactor during experiment G12. Full symbols reactor: heated to 400 °C. Open symbols: reactor not actively heated (heating switched off). G16-W: temperature profile of water before experiment G16 (same temperature set points of the preheater, the salt separator and the reactor and same mass flow conditions).

bed. This behavior was observed for all gasification runs in the presence of K3PO4 compared to the respective run without salt, e.g., G14a and G14b or G16a and G16b (run a refers to the gasification without K3PO4, run b refers to the gasification with K3PO4, see also the Supporting Information). However, this behavior was not observed for experiment G17 (refer to the Supporting Information), which might be explained by a much lower reactor inlet temperature of about 380 °C in that experiment and by the superposition of the reactor heater control. 3.3. Influence of K3PO4 on the Catalytic Gasification of Glycerol. When gasifying the mixture of 20 wt % glycerol with 0.05 mol/kg K3PO4 (runs G12c-d), the salt breakthrough (increase in the conductivity of the reactor effluent) was 8412



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catalysts with a certain amount of Li, Na, K, or Cs enhanced the water gas shift reaction significantly. This effect was described as “permanent strong metal-support interactions” caused by the interaction of the alkali metals with the support. On the other hand, alkali metal dopants promote the formation of higher hydrocarbons in the hydrogenation of CO over ruthenium catalysts (Fischer−Tropsch synthesis).72,73,75,75 For example, potassium decreases the rate of formation for all products but methane is more affected; thereby the selectivity toward higher alkanes and alkenes is increased.72,73 However, no information was found on the effect of alkali metals or phosphorus on the gasification performance of ruthenium catalysts. Perhaps the steam reforming activity of the Ru/C catalyst was lowered in the presence of the salt in the reactor. Thus, the higher hydrocarbons would result from an incomplete degradation of the glycerol (see the discussion above). The formation of the higher hydrocarbons via Fischer−Tropsch synthesis can be ruled out because in that case higher hydrocarbons should also have been observed in the gasification of glycerol without K3PO4. In other words, if Fischer−Tropsch synthesis had played a considerable role higher hydrocarbons would have always been present in the product gas. The presence of potassium (in K3PO4) would only have led to a changed selectivity of methane to higher hydrocarbons. 3.4. Some Comments on Corrosion. In this study we detected Fe as the only corrosion product in the brine effluent and in the reactor inlet but not in the reactor effluent. However, the Fe concentration was close to the detection limit of the ICP-OES instrument due to the high dilution of the sample. Since we focused on a quantitative analysis of potassium and phosphorus the samples had to be diluted by factors of 12.5− 500 to match the calibrated range. However, corrosion products (Fe, Cr, Ni, and Ti) were found on the catalyst that was removed after this study.22,71 Apparently, all corrosion products had deposited on the catalyst.

Article

ASSOCIATED CONTENT

S Supporting Information *

The Supporting Information contains: (1) The detailed experimental parameters of this study, (2) a list of chemicals used in this study, (3) a detailed description of all applied analytical methods, (4) densities of representative solutions of the brine and the reactor effluent, (5) assumptions and equations for the estimation of the uncertainties of the mass balance, carbon gasification efficiency and the methane yield, (6) representative temperature profiles of the fluid inside the preheater, and (7) detailed results regarding the carbon gasification efficiency, the methane yield, and the material flows of all effluents for all gasification experiments. This material is available free of charge via the Internet at http:// pubs.acs.org.

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AUTHOR INFORMATION

Corresponding Author

*Phone: +41 (0)56 310 2135. Fax +41 (0)56 310 2199. E-mail: [email protected]. Present Address †

Karlsruhe Institute of Technology, D-76131 Karlsruhe, Germany. Notes

The authors declare no competing financial interest.

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ACKNOWLEDGMENTS The authors would like to thank Axpo Naturstrom-Fonds and the Competence Center for Energy and Mobility in Switzerland (CCEM-CH) for financial support. Further, the authors gratefully acknowledge E. De Boni, T.-B. Truong, A. Schuler, J. Schneebeli, and V. Sette Gripp for experimental support and M. Brandenberger, H. Zöhrer, S. Stucki, T. Schildhauer, J. Wambach, C. Ludwig, A. Wokaun and M. Mazzotti for valuable discussions.

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4. SUMMARY AND CONCLUSIONS The principal concept of the hydrothermal gasification process developed at PSI, that is the continuous gasification with the simultaneous salt separation, was successfully shown for the gasification of solutions of pure glycerol with K3PO4. The following conclusions can be drawn from results presented in this article: •Glycerol can be gasified efficiently to a methane-rich gas. The gas composition corresponds to the thermodynamic equilibrium. Neither tars nor char were formed in the catalytic gasification and methanation. •The simultaneous gasification and methanation can be operated in an autothermal mode in a well-insulated fixed-bed reactor. •The presence of salts, e.g., potassium phosphates, in the catalytic reactor leads to a shift in the gas composition toward C2−C4 hydrocarbons. The higher hydrocarbons might result from incomplete gasification of the glycerol in the presence of potassium and phosphate in the reactor. This effect is reversible. Thus, neither potassium nor phosphate can be considered as poisons for the Ru/C catalyst. Note, however, that higher hydrocarbons can also be formed without the presence of salt when the weight hourly space velocity is too high for complete gasification of the glycerol. Thus, the effect of the K3PO4 is likely a physical blockage of a part of the catalyst bed, which is equivalent to a higher effective WHSV.

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Continuous Hydrothermal Gasification of Glycerol Mixtures

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