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2013 122 Gorka García Sáinz Estudio experimental de la cogasificación de fangos de EDAR y carbón con aire en lecho fluidizado Director/es Departamento Ingeniería Química y Tecnologías del Medio Ambiente Ábrego Garrués, Javier Sánchez Cebrián, José Luis Director/es Tesis Doctoral Autor Repositorio de la Universidad de Zaragoza – Zaguan http://zaguan.unizar.es UNIVERSIDAD DE ZARAGOZA Departamento Director/es Gorka García Sáinz ESTUDIO EXPERIMENTAL DE LA COGASIFICACIÓN DE FANGOS DE EDAR Y CARBÓN CON AIRE EN LECHO FLUIDIZADO Director/es Ingeniería Química y Tecnologías del Medio Ambiente Ábrego Garrués, Javier Sánchez Cebrián, José Luis Tesis Doctoral Autor 2013 Repositorio de la Universidad de Zaragoza – Zaguan http://zaguan.unizar.es UNIVERSIDAD DE ZARAGOZA Departamento Director/es Director/es Tesis Doctoral Autor Repositorio de la Universidad de Zaragoza – Zaguan http://zaguan.unizar.es UNIVERSIDAD DE ZARAGOZA TESIS DOCTORAL Universidad de Zaragoza Departamento de Ingeniería Química y Tecnologías del Medio Ambiente Estudio experimental de la co-gasificación de fangos de EDAR y carbón con aire en lecho fluidizado Septiembre de 2013 Gorka García Sáinz TESIS DOCTORAL Universidad de Zaragoza Departamento de Ingeniería Química y Tecnologías del Medio Ambiente Estudio experimental de la co-gasificación de fangos de EDAR y carbón con aire en lecho fluidizado Memoria que presenta para optar al Grado de Doctor en Ingeniería Química Gorka García Sáinz Zaragoza, Septiembre de 2013 La presente Tesis Doctoral, que lleva por título “Estudio experimental de la co-gasificación de fangos de EDAR y carbón con aire en lecho fluidizado”, realizada por D. Gorka García Sainz y dirigida por los doctores D. Javier Ábrego Garrués y D. José Luis Sánchez Cebrián se presenta como compendio de las siguientes publicaciones: I. G. García, J. Arauzo, A. Gonzalo, J.L. Sánchez, J. Ábrego, Influence of feedstock composition in fluidised bed co-gasification of mixtures of lignite, bituminous coal and sewage sludge, Chem. Eng. J., 222 (2013) 345–352. II. G. García, E. Cascarosa, J. Ábrego, A. Gonzalo, J.L. Sánchez, Use of different residues for high temperature desulphurisation of gasification gas, Chem. Eng. J., 174 (2011) 644-651. III. G. García, A. Monzón, F. Bimbela, J.L. Sánchez, J. Ábrego, Desulfurization and catalytic gas cleaning in fluidized bed co-gasification of sewage sludge-coal blends, Energy Fuels, en prensa (2013). DOI: 10.1021/ef400259g IV. G. García, E. Campos, I. Fonts, J.L. Sánchez, J. Herguido, Gas Catalytic Upgrading In A Two Zone Fluidized Bed Reactor Coupled To A Co-Gasification Plant, Energy Fuels, en prensa (2013). DOI: 10.1021/ef400227z Lo que se hace constar en cumplimiento del Reglamento sobre Tesis Doctorales de la Universidad de Zaragoza (artículo 17, apartado a), aprobado según el acuerdo de 17 de diciembre de 2008 del Consejo de Gobierno de la Universidad gasificador. En este caso el estudio está justificado porque tanto el lodo de EDAR como los carbones que se alimentan tienen un elevado contenido en cenizas, que pueden tener un efecto catalítico en el propio gasificador donde se introducen [11]. Este estudio constituyó el objeto del artículo I (Influence of feedstock composition in fluidised bed co-gasification of mixtures of lignite, bituminous coal and sewage sludge, publicado en la revista Chemical Engineering Journal), donde el efecto de la composición de la alimentación se estudió mediante un diseño experimental de mezclas, analizándose los resultados estadísticamente mediante Análisis de Varianza (ANOVA). El gas que se obtiene en la gasificación contiene diversos contaminantes de distinta naturaleza [12]. Para algunos de ellos, como las partículas, existe tecnología industrialmente conocida y fiable para su eliminación. Sin embargo, la presencia de otros compuestos -como son los alquitranes, hidrocarburos condensables que se forman durante la etapa de descomposición térmica o pirólisis de los materiales carbonososexige el desarrollo de nuevos sistemas de limpieza, dado que a escala industrial, y hasta la fecha, su eliminación no se ha resuelto de manera satisfactoria. Es por ello que una buena parte de la investigación que se realiza actualmente en el proceso de gasificación se centra en la eliminación de estos compuestos del gas que se produce [13–15]. En el caso de la co-gasificación de lodos y los carbones seleccionados en este trabajo, se encuentra una dificultad adicional, la presencia de cantidades significativas de azufre en su composición, que conlleva la formación de gases sulfurados, siendo el sulfuro de hidrógeno, H2S, el más importante [16]. La presencia de este gas implica que los catalizadores más ampliamente usados para el craqueo de hidrocarburos, basados en Ni, no pueden usarse directamente, dado que el H2S es un veneno que desactiva el Ni por formación de NiS de forma muy rápida e irreversible [17]. 13 Es por lo expuesto en el párrafo anterior que se planteó un estudio de la desulfuración del gas a elevada temperatura, para que se pudiera aplicar con el gas tal y como se obtiene en el gasificador. Este aspecto viene condicionado por la presencia de alquitranes, que condensan a temperaturas inferiores a 300-400 ºC, causando graves problemas en tuberías y filtros, por ejemplo [15]. Además, el tratamiento del gas en caliente tiene la ventaja de mejorar la eficiencia energética del proceso, dado que no se pierde energía del gas asociada a su calor sensible. Este estudio se realizó con gas sintético, que simula el gas producido en la co-gasificación de carbón y lodo. En esta parte de la Tesis Doctoral se usaron como materiales desulfurantes las propias cenizas y char obtenidas de los tres materiales a gasificar (lodo, hulla y lignito). Para comparar los resultados obtenidos con estos materiales, se usó dolomita calcinada, que es un material comercial de bajo precio, previamente estudiada por diversos autores en la desulfuración de gases de distinto origen, entre ellos el procedente de la gasificación de biomasa y carbón [18–21]. El trabajo realizado fue publicado en el artículo II (Use of different residues for high temperature desulphurisation of gasification gas, publicado en la revista Chemical Engineering Journal). Si en los primeros trabajos publicados dentro de esta Tesis Doctoral se fijaron la composición de la alimentación más adecuada y las condiciones para eliminar el H2S que se produce en la gasificación, lo siguientes trabajos abordan la limpieza del gas y mejora de sus características, con el objetivo de reducir el alquitrán presente, lo que permitiría el uso del gas para la producción de electricidad en un motor de combustión interna o una turbina. La primera estrategia que se estudió fue más convencional, estudiando la limpieza catalítica del gas usando un catalizador comercial que, por su tamaño de partícula, debe ser usado en lecho fijo. Así en este trabajo, que se corresponde en el artículo III (Desulfurization and catalytic gas cleaning in fluidized bed co-gasification of sewage 14 sludge-coal blends, publicado en la revista Energy and Fuels) se estudió el efecto de la temperatura del lecho catalítico en la calidad del gas obtenido. Por último, se planteó el uso de un reactor novedoso [22], desarrollado por el Grupo de Investigación de Catálisis, Separaciones Moleculares e Ingeniería del Reactor (CREG), perteneciente también al Departamento de Ingeniería Química y Tecnologías del Medio Ambiente de la Universidad de Zaragoza, que no ha sido usado previamente para la limpieza y mejora del gas de gasificación. Se trata del Reactor de Lecho Fluidizado de Dos Zonas (RLFDZ o TZFBR en inglés), reactor que permite integrar en un solo equipo el reformado del gas y la regeneración del catalizador ante la desactivación por deposición de coque. En este trabajo, que se corresponde con el artículo IV (Gas Catalytic Upgrading In A Two Zone Fluidized Bed Reactor Coupled To A Co-Gasification Plant, publicado en la revista Energy and Fuels), se estudió el efecto de la concentración de oxígeno alimentada en la zona de regeneración del reactor sobre el rendimiento y la calidad de los productos obtenidos. 15 16 2. Resumen 2.1. Objetivos Como se ha comentado en el Apartado 1 de esta Memoria, el estudio que ha llevado a la realización de esta Tesis Doctoral surge del proyecto “Coal Catalytic Co-Gasification in an Innovative Rotary Kiln Gasifier” (COCACORK), que tenía como objetivo el estudio y mejora del aprovechamiento de carbón y biomasa mediante gasificación. La gasificación de biomasa puede permitir la producción de energía mediante fuentes localmente disponibles, disminuyendo la dependencia energética de algunos países, como es el caso de España. El uso de carbón local también puede ser económica y estratégicamente beneficioso, aunque en algunos países, el carbón del que se dispone es de baja calidad, en concreto por la presencia de compuestos como el azufre que supone una contaminación atmosférica en caso de emitirse con los humos de la combustión. Este es también el caso de España, donde algunos de los carbones que se pueden extraer tienen un elevado contenido en azufre que los hace no aptos para su combustión en centrales térmicas de carbón, o al menos impone la necesidad de alimentarlos mezclados con otro carbón de mejores características. En este escenario, se planteó como objetivo general el estudio de la co-gasificación de fango de EDAR en un reactor de lecho fluidizado y usando aire como agente gasificante, dando especial importancia a la limpieza del gas y mejora de sus características. Como objetivos parciales se establecieron los siguientes: •Estudio de las condiciones de operación e influencia sobre los productos obtenidos. •Estudio de la eliminación del H2S del gas de gasificación 17 •Estudio de la limpieza de alquitranes del gas. 2.2. Antecedentes La gasificación es el proceso termoquímico por el cual un sustrato carbonoso se transforma mediante una oxidación parcial en un gas combustible que contiene, entre otros compuestos, hidrógeno, monóxido de carbono, metano, dióxido de carbono, y nitrógeno, así como diversas impurezas o contaminantes. La conversión a un gas combustible permite su uso en aplicaciones distintas a la combustión convencional para generar calor o vapor para un ciclo Rankine: el gas puede llevarse a combustión en motores de combustión interna o turbinas para producción de electricidad o cogeneración simultánea de electricidad y calor, o usarse para la síntesis de productos de mayor valor añadido, como metanol u otros alcoholes, o hidrocarburos con características del diésel, la gasolina o ceras, mediante la síntesis Fischer-Tropsch. La gasificación de biomasa se contempla actualmente como una tecnología que puede contribuir a disminuir la dependencia de los combustibles fósiles y ampliar el espectro de opciones para la producción de energía, aumentando el peso que la biomasa tiene dentro de las energías primarias que actualmente se utilizan. Dentro de la amplitud de tipos distintos de biomasa que se encuentran, el Grupo de Investigación de Procesos Termoquímicos se encuentra trabajando activamente en la aplicación de los procesos de pirólisis y gasificación de fango de EDAR. El fango de EDAR se obtiene como resultado del proceso de tratamiento biológico de aguas residuales, y es un residuo cuya generación ha aumentado considerablemente debido a la aplicación de una legislación medioambiental cada vez más estricta [23]. La co-gasificación, que consiste en alimentar una mezcla de sustratos al reactor de gasificación, presenta una serie de ventajas, como es por una parte el aumento de 18 escala, ya que cuando se trata de biomasa, la disponibilidad de la misma es un aspecto crítico a la hora de escalar el proceso, y por tanto de su viabilidad económica. En el caso de usar carbón, esto implica que la disponibilidad de biomasa no es un aspecto tan crucial. Por otra parte, los procesos de gasificación de carbón a gran escala han despertado históricamente un mayor interés que los de biomasa o residuos, por lo que la co-gasificación con fango a escala industrial se beneficiaría del mayor desarrollo de la tecnología de gasificación de carbón, a la vez que permitiría el aprovechamiento energético de un residuo que se produce en cantidades ingentes y cuyo destino final es problemático. Uno de los aspectos clave del proceso de gasificación es la formación de alquitranes, compuestos orgánicos condensables que se forman durante la gasificación mediante una serie de reacciones complejas y fuertemente dependientes de las condiciones de operación. Conforme la temperatura de reacción se incrementa, se va produciendo un “esquema de formación de alquitranes” [24], mostrado en la Figura 1. Figura 1. Esquema de formación de alquitranes. Además de la temperatura, la presión y el tiempo de residencia también influyen en el tipo de alquitrán obtenido. En la Tabla 1 se muestra una clasificación más detallada. Existen distintas estrategias para eliminar estos compuestos del gas, que se pueden dividir en sistemas que buscan separar el alquitrán del gas, mediante lavado o filtrado, o sistemas basados en la destrucción o craqueo de los alquitranes. La ventaja obvia de estos últimos es que no se genera una corriente que debe ser tratada a posteriori. Mezcla de productos oxigenados Éteres fenólicos Éteres alquílicos Éteres heterocíclicos Hidrocarburos poliaromáticos (PAH) PAH de alto peso molecular 400ºC 500ºC 600ºC 700ºC 900ºC 800ºC 19 Además, el craqueo de alquitranes permite enriquecer el gas de gasificación, ya que se forman hidrocarburos más ligeros, a la vez que aumenta la eficiencia de conversión de la materia prima gasificada. Tabla 1. Clasificación y propiedades de los alquitranes según Li y Suzuki [24]. Clase Nombre Propiedades Compuestos representativos 1 Indetectables por Cromatografía de Gases (CG) Muy pesados, incapaces de ser detectados por CG. Se determinan por diferencia con el resto de los alquitranes detectables. 2 Aromáticos heterocíclicos Contienen heteroátomos, altamente solubles en agua Piridina, fenol, cresoles, quinolina, isoquinolina, dibenzofenol 3 Aromáticos ligeros (un anillo) No suponen ningún problema en cuanto a su condensabilidad y solubilidad. Tolueno, etilbenceno, xilenos, estireno 4 PAHs ligeros (2-3 anillos) Condensan a bajas temperaturas incluso a concentraciones muy bajas. Indeno, naftaleno, metilnaftaleno, fluoreno, antraceno. 5 PAHs pesados (4-7 anillos) Condensan a altas temperaturas y bajas concentraciones. Fluoranteno, pireno, criseno, perileno, coroneno. En el craqueo se pueden diferenciar: •Craqueo térmico: la ruptura se produce por medio de exposición a temperaturas elevadas (de 900 a 1200 ºC). Sin embargo es difícil conseguir un craqueo completo, y deben afrontarse problemas operacionales, inversiones económicas a veces excesivas, y aportes energéticos adicionales. •Craqueo catalítico: implica el uso de un catalizador y facilita la selectividad a determinados gases. Es una técnica que ha despertado un amplio interés. En el siguiente apartado se tratará este método en mayor profundidad. 20 Las reacciones implicadas en la descomposición de los alquitranes han sido ampliamente descritas por diversos autores [25,26], siendo las principales las mostradas en la Tabla 2. Tabla 2. Reacciones principales en el craqueo de alquitranes. Reacción Ecuación estequiométrica Craqueo Térmico pC n H x  qC m H y + rH 2 (m<n) Steam Reforming C n H m + nH 2 O nCO + (n+m/2)H 2 Dry Reforming C n H m + nCO 2 2nCO + (m/2)H 2 Formación de Carbono C n H m  nC + (m/2)H 2 CH 4  C + 2H 2 El craqueo catalítico permite una mejor eficacia energética, dado que se opera a menor temperatura, y mayor especificidad en la composición del gas que se obtiene. Entre los catalizadores más usados, destacan los basados en óxidos de níquel, que han sido ampliamente estudiados para mejorar el gas procedente de la gasificación de biomasa [13,27]. Además, tanto el fango de EDAR como algunos carbones contienen cantidades significativas de azufre, que en el proceso de gasificación conducen a la generación de compuestos indeseados como el sulfuro de hidrógeno. Por tanto, se hace necesario plantear otra estrategia adicional para la limpieza y mejora de la calidad del gas. La eliminación del sulfuro de hidrógeno que se forma durante la gasificación es crítica para el éxito de la tecnología de gasificación de estos materiales que se usan en la Tesis Doctoral, especialmente si se considera el alto contenido en azufre de partida de las materias primas alimentadas, como el fango de EDAR o el lignito turolense. Además de evitar problemas con la lluvia ácida [28], se puede producir corrosión en las 21 tuberías, turbinas y otros equipos. Además, si hay H2S en el gas, el uso de catalizadores basados en Ni (metal clásico en el reformado de hidrocarburos) se ve seriamente limitado por el envenenamiento debido a la formación de NiS [13,17]. La eliminación del H2S se puede realizar mediante procesos a baja temperatura, como el lavado con disoluciones acuosas, o por procesos de desulfuración a alta temperatura, que presentan la ventaja de una mayor eficiencia energética y se pueden aplicar tanto en el mismo lecho fluidizado como tras el gasificador. Los procesos a alta temperatura han sido revisados recientemente en el trabajo de Meng y cols. [29]. Aunque se han probado distintos compuestos metálicos, los compuestos con calcio son los más usados en procesos de gasificación y co-gasificación [20,30]. Entre los compuestos con calcio, la dolomita es uno de los más eficaces, alcanzado rendimientos en la desulfuración superiores al 90% [31,32]. Algunos autores han usado la dolomita para evitar la desactivación de un catalizador de Ni con resultados prometedores [28,33,34]. El principal problema asociado a este material es la facilidad con que se produce atrición en lechos fluidizados, limitando su aplicación en este tipo de reactores para procesos de gasificación y co-gasificación. 22 controlador de flujo másico, usándose un caudal de 2.81 LN/min. El caudal de alimentación se regula mediante un variador de frecuencia conectado al motor de un tornillo sin fin. En los distintos experimentos realizados se usó una relación estequiométrica o de equivalencia del 30 % (aire alimentado respecto al necesario para la combustión completa) variando el caudal de sólido en función de su composición. La temperatura del reactor se fijó en 850ºC, al igual que la temperatura del freeboard, mediante un horno eléctrico que permite un control independiente de la temperatura de las distintas zonas. El gas que sale del reactor pasa por un ciclón y un filtro en caliente, ambos a una temperatura de 450ºC. Inicialmente en el reactor se introducen 300 g de arena con un tamaño de partícula de 250-350 µm. Los experimentos realizados han tenido una duración de 90 min una vez comenzada la alimentación, tiempo suficiente para alcanzar unas condiciones estables en el reactor y obtener una distribución y composición de los productos representativa. A la salida del filtro en caliente, el gas se enfría mediante dos condensadores de vidrio refrigerados con hielo y pasa por un filtro de algodón que sirve para retener la niebla de alquitranes que no ha precipitado. Mediante un cromatógrafo de gases (Agilent Micro-GC 3000) se analiza el gas cada 5 min aproximadamente. Al finalizar el experimento se pesan los sólidos recogidos del lecho y ciclón, mientras los aparatos de vidrio se lavan con isopropanol. El agua condensada me mide mediante valoración Karl-Fischer y el alquitrán se calcula gravimétricamente por diferencia con el agua. Reactor de lecho fijo para el estudio de desulfuración del gas. En la Figura 3 se muestra un esquema del sistema experimental usado. El reactor está fabricado en cuarzo, con un diámetro interno de 1.2 cm y 40 cm de altura. El sólido, alrededor de 1 g en cada experimentos, se coloca sobre fibra de vidrio, que actúa de soporte, a 18.5 cm de la parte superior del reactor, en la zona isoterma del mismo. Una vez situado el sólido, el reactor se calienta mediante un horno eléctrico hasta la temperatura final (700, 800 o 900ºC) bajo un flujo de nitrógeno. Una vez alcanzada la temperatura, se 29 introduce el gas a desulfurar, con un caudal de 0.05 LN/min, mediante un controlador de flujo másico, fijándose en los experimentos un tiempo espacial (WHSP, weight hourly space time) de 3.6 h-1. La composición del gas tras el reactor se analiza con un cromatógrafo Agilent Micro-GC 3000. Figura 3. Reactor usado para el estudio de desulfuración Reactor de lecho fijo usado con catalizador de Ni comercial. Este reactor se situó en serie con el de lecho fluidizado ya descrito, tal y como se muestra en la Figura 4. La conexión entre ambos reactores se hace mediante una tubería calentada externamente con una resistencia eléctrica, manteniendo su temperatura a 450ºC para evitar la condensación del alquitrán. El reactor de lecho fijo usado tiene un diámetro interno de 35.5 mm y 990 mm de longitud. Se encuentra calentado mediante un horno eléctrico con dos zonas independientes de calentamiento, lo que permite fijar temperaturas distintas en la parte superior e inferior. El gas, que entra por la parte inferior del reactor, se encuentra en primer lugar con un lecho de dolomita, 30 previamente calcinada (250 g de dolomita fresca, que resultan en aproximadamente 133 g tras la calcinación). F T T 1 9 T 4 T T TT 11 F 2 4 44 4 4 4 5 6 F 78 6a 6c 9 10 2a 2b 12 13 14 15 16 P-12 1 Reactor de lecho fluidizado 2 Reactor de lecho fijo 2a Lecho de dolomita 2b Lecho de Ni 3 Horno eléctrico 4 Controladores de temperatura 5 Tolva y tornillo sinfín 6 Compresor de aire 6a Entrada aire refrigeración 6b Salida aire refrigeración 6c Entrada aire gasificación 7 Hidrógeno 8 Nitrógeno 9 Ciclón 10 Depósito de char 11 Filtro alta temperatura 12 Condensadores 13 Filtro de algodón 14 Contador de gas 15 MicroGC 16 Salida de gas 33 6b Figura 4. Sistema experimental usado en la limpieza de gas con catalizador comercial Encima del lecho de dolomita se sitúa el catalizador de níquel (135 g), en forma de monolitos. Entre el catalizador y la dolomita se coloca fibra de cuarzo, con el fin de evitar la fluidización de la dolomita, ya que se trata de un material que se desmenuza fácilmente. Antes de cada experimentos, el catalizador de Ni se redujo a 800ºC durante dos horas, con un flujo de 2. 82 LN/min de H2 (15 % en N2). La temperatura de la dolomita, que se seleccionó como material desulfurante, se fijó en 800 ºC, mientras que el catalizador de Ni se usó a 800, 850 y 900ºC. Reactor de lecho fluidizado de dos zonas. Para el estudio de limpieza del gas con el reactor de lecho fluidizado de dos zonas se usaron tanto el reactor de gasificación como el reactor de lecho fijo, en este caso únicamente con dolomita para desulfurar el gas. Tras el reactor de lecho fijo se tomó una parte del gas producido, 31 aproximadamente un 10 % del caudal, para su tratamiento en el reactor de dos zonas, tal y como se muestra en la Figura 5. Figura 5. Sistema experimental usado en la limpieza del gas con el RLFDZ. Tras el reactor de desulfuración donde se encuentra la dolomita, el gas se divide en dos corrientes. La mayor parte (unos 3 LN/min) pasan a un sistema de condensación y análisis, como ya se ha descrito anteriormente. El resto del gas (0.3 LN/min) se dirige al reactor de lecho fluidizado de dos zonas (RLFDZ), que se muestra en la Figura 6. En este reactor es posible llevar a cabo las dos etapas típicas de un proceso catalítico, reacción y regeneración del catalizador. El reactor está fabricado en cuarzo, con una altura total de 510 mm. La zona de regeneración (inferior) tiene un diámetro de 18 mm, mientras que la zona de reformado tiene 28 mm de diámetro interno. El gas procedente de reactor de desulfuración entra al RLFDZ por la parte superior, a través de un tubo de cuarzo. 32 Figura 6. RLFDZ El gas para la regeneración (O2 y N2 en distinta proporción) entra por la parte inferior, de modo que el movimiento del catalizador (42 g de Ni/γ-Al2O3) entre las dos zonas permite que realice el reformado y se regenere por combustión del coque depositado en continuo. En cada instante, aproximadamente 17 g de catalizador están en la zona de regeneración. El uso de una cantidad adecuada de O2, permite que no se produzca combustión del gas procedente de la gasificación. Un parámetro de especial importancia para el funcionamiento de este reactor es la relación u/umf (velocidad superficial del gas/velocidad de mínima fluidización) en las dos secciones del lecho [22]. En la experimentación llevada a cabo, se ha usado un caudal de la mezcla oxígeno/nitrógeno de regeneración de 0.3 LN/min, para obtener 33 una u/umf = 2.5 a la temperatura de operación. En la zona de reformado la u/umf es de 3.4, calculada teniendo en cuenta el cambio de dimensiones y el flujo de gas proveniente de la zona de regeneración junto con el caudal de gas a reformar, que como se ha comentado es de 0.3 LN/min. Hay que tener en cuenta que el uso de nitrógeno, que se eligió por simplicidad experimental, implica una dilución del gas reformado, dilución que se ha descontado en los cálculos. En el caso de una operación a escala industrial se debería usar otro compuesto, como vapor de agua, que fuera fácilmente eliminable, o usar un RLFDZ con dimensiones apropiadas para usar oxígeno puro, por ejemplo. Antes de comenzar el experimento, se llevó a cabo la reducción del catalizador con H2 (5%, diluido en nitrógeno) a 800ºC. La temperatura en el RLFDZ se fijó también en 800ºC mediante un horno eléctrico. 34 2.4. Conclusiones y trabajos futuros En este Apartado de la Memoria se van a mostrar de forma resumida la experimentación realizada y las conclusiones más relevantes obtenidas en cada una de las partes de la que se compone este trabajo. Para mayor claridad, se ha estructurado siguiendo cada uno de los trabajos publicados, añadiendo además un apartado donde se consideran posibles trabajos de investigación que continuarían la línea de investigación que supone esta Tesis Doctoral. 2.4.1. Influencia de la composición de la alimentación El objetivo de este estudio fue determinar cuál era el efecto de modificar la composición de una mezcla de carbón bituminoso (B), lignito (L) y fango de EDAR (SS) en los productos obtenidos en la gasificación. Estos materiales tienen un importante contenido en ceniza, en cuya composición se encuentran distintos metales, mostrados en la Tabla 5, como Al, Fe, Ca, Mg, K y Na, que pueden tener un efecto catalítico durante la gasificación [2,35–38]. Para ello se usó el reactor de lecho fluidizado descrito anteriormente, planteando para el estudio un diseño de experimentos de mezclas. En la Figura 7 se muestra, sobre un diagrama ternario, la composición de las distintas mezclas utilizadas. 35 0.00 0.25 0.50 0.75 1.00 0.00 0.25 0.50 0.75 1.00 0.00 0.25 0.50 0.75 1.00 Lignito Fango EDAR Hulla Figura 7. Mezclas usadas en el diseño experimental En la Tabla 8 se muestran algunos de los resultados experimentales obtenidos, en concreto la concentración de alquitrán en el gas (g/m3N), el % de alquitrán obtenido frente al sólido alimentado (%), el % de carbono alimentado que se convierte en gas (ηC, %), el rendimiento energético a gas (ηenerg, %) calculado según la Ecuación 1, el poder calorífico inferior del gas seco obtenido (kJ/m3N) y la producción específica de gas (Ygas, m3N/kg). En la Tabla 9 se muestra la composición media del gas obtenido. 𝜂𝑒𝑛𝑒𝑟𝑔 (%)=𝑃𝐶𝐼𝑔𝑎𝑠 ·𝑌 𝑔𝑎𝑠 𝑃𝐶𝐼𝑓𝑎𝑛𝑔𝑜 ·100 Ec. 1 El análisis de estos resultados se realizó mediante ANOVA (análisis de varianza), que se basa en comparar la varianza experimental asociada al error (que se calcula a partir de las repeticiones que se realizan de uno o varios experimentos) con la varianza que crea, en este caso, la variación de la composición. La comparación se realiza mediante un test F de Fischer, y permite discriminar si el efecto observado es estadísticamente 36 significativo frente al error, con un nivel de confianza predeterminado (95% en este estudio). Tabla 8. Resultados experimentales (1). Alquitrán, rendimiento de carbono a gas y energético, PCI del gas y producción específica de gas Fracción masa B/L/SS Alquitrán g/m3N Alquitrán % η C (%) η energ (%) PCI kJ/m3(STP) Y gas m3N/kg 1/0/0 41.2/43.8 12.3/12.6 42.1/42.6 23.2/22.6 2293/2201 3.00/3.13 0/1/0 58.4/64.5 15.6/15.6 47.2/46.6 20.2/22.4 1686/2070 2.67/2.42 0/0/1 38.4/64.2 10.2/17.3 79.7/79.5 47.3/51.7 4112/3949 2.65/2.69 0.5/0.5/0 41 11.7 52 22.7 2044 2.86 0.5/0/0.5 21.9/28.7 6.1/7.7 56.0/55.0 35.0/33.2 3191/3103 2.81/2.69 0/0.5/0.5 23.0/18.2 6.0/4.9 64.6/66.7 40.8/42.7 3293/3322 2.62/2.71 0.67/0.17/0.17 28.4 8.3 48.6 24.3 2302 2.9 0.17/0.67/0.17 5.2 1.4 52.7 25 2071 2.77 0.17/0.17/0.67 38.1 10.2 64.3 40.5 3477 2.69 0.41/0.41/0.18 12.5 3.1 43.3 20.5 2090 2.46 0.33/0.33/0.33 6.4/12.9 20.5/20.7 1.6/3.5 5.4/5.5 50.1/51.3 46.9/46.4 29.7/30.5 28.4/27.9 2882/2731 2585/2510 2.51/2.7 2.62/2.66 Tabla 9. Resultados experimentales (2). Composición media del gas (% vol.) B/L/SS H2 CO CH4 C2Hn H2S 1/0/0 5.73/5.09 9.04/8.90 1.15/1.16 0.20/0.23 0.01/0.04 0/1/0 5.44/7.50 6.01/7.84 0.80/0.60 0.09/0.10 1.00/0.82 0/0/1 8.24/7.56 8.52/8.14 3.13/3.14 1.75/1.67 0.10/0.11 0.5/0.5/0 6.23 7.73 0.92 0.11 0.2 0.5/0/0.5 8.20/8.09 9.31/9.11 1.97/1.87 0.72/0.69 0.04/0.05 0/0.5/0.5 8.87/9.04 9.26/9.53 2.23/2.00 0.74/0.72 0.41/0.42 0.67/0.17/0.17 6.41 8.25 1.13 0.29 0.07 0.17/0.67/0.17 6.55 7.46 0.86 0.2 0.31 0.17/0.17/0.67 8.81 9.32 2.15 0.99 0.14 0.41/0.41/0.18 5.9 8.07 0.89 0.2 0.11 0.33/0.33/0.33 8.10/7.55 6.86/6.65 9.24/8.87 9.77/9.69 1.54/1.45 1.10/1.01 0.49/0.47 0.37/0.35 0.13/0.32 0.14/0.12 A partir de los efectos significativos encontrados con el ANOVA, es posible modelar empíricamente el efecto sobre las distintas variables medidas con una ecuación como la que se muestra en la Ecuación 2: 37 RV= β 1·W1+ β 2·W2+ β 3·W3 +β 12·W1·W2 +β 13·W1·W3 +β 23·W2·W3 +β 123·W1·W2·W3 Ec. 2 Donde RV es cualquiera de los resultados experimentales mostrados, los coeficientes β i representan la respuesta lineal frente a la fracción másica (Wi) de los materiales puros, los coeficientes β ij son el efecto cuadrático debido al efecto sinérgico o antagónico debido a la mezcla de dos de los materiales y el coeficiente β 123 modela el efecto cúbico debido a la mezcla de los tres materiales. Cuando un efecto no es significativo, su coeficiente no aparece en la ecuación. En la Tabla 10 se muestran los coeficientes de regresión obtenidos tras el ANOVA, así como el valor del coeficiente de correlación R2 y R2 ajustado. Estos dos últimos coeficientes dan una idea de cómo se ajustan los datos experimentales al modelo empírico. 38 La comparación de estos datos con los datos de referencia correspondientes a los experimentos en blanco permite el cálculo de la eficiencia de desulfuración de cada uno de los materiales a lo largo del tiempo. La Tabla 11 presenta un resumen de la eficiencia promedio (% de S capturado por el lecho) y de la capacidad de eliminación del H2S (mg de S por gramo de material de lecho a los 100 minutos de experimento), para todos los materiales empleados en esta sección. Como puede observarse en la Tabla 11, el char de lignito, hulla y fango empleado tiene una capacidad de desulfuración modesta, hecho ya evidenciado en la Figura 11. Resultan llamativos los valores negativos de eficiencia del char de lignito a 800 y 900ºC, lo que puede atribuirse al alto contenido original en S de este material (que puede producir liberación adicional de H2S a altas temperaturas). Por su parte, el char de fango ofrece mejores resultados, siendo capaz de capturar algo más del 50% del S presente en la alimentación; expresado en términos de capacidad, puede retener 12.1 mg S/g de material. Este valor es similar a los encontrados en la bibliografía disponible [40,41]. Tabla 11. Eficiencia y capacidad de desulfuración de los materiales empleados. Eficiencia promedio de desulfuración (% de S eliminado del gas) Capacidad de desulfuración (tras 100 min) (mg de S por gramo of sólido) T(ºC) 700 °C 800 °C 900 °C 700 °C 800 °C 900 °C BC 17.8 15.8 23 4.7 3.1 5.1 LC 10.1 −39.7 −15.6 2.8 −1.7 −2.0 SSC 18.9 35.3 50.5 4.5 10.4 12.1 BA 72.2 97.2 97.6 23.2 24 21.3 LA 97.7 94.9 92.5 26.7 23.9 20.1 SSA 34.8 46.4 29.7 10.6 13.9 7.4 DO 94.5 99.3 98.4 26.8 24.8 21.5 45 Atendiendo a la composición y superficie BET del char de los tres materiales, mostrados en la Tabla 4, puede concluirse que su capacidad de desulfuración es independiente de la superficie específica de los mismos, y puede estar directamente relacionada con el contenido en cenizas y en especial de ciertos óxidos de metales, tales como Fe y Ca (este último muy abundante en el fango) y su capacidad para reaccionar directamente con el H2S a altas temperaturas [5] según la Ecuación 3: MxOy +y H2S → MxSy + y H2O Ec. 3 En el caso de las cenizas, los resultados son notablemente diferentes. Para lignito y hulla (LA y BA), las eficiencias y capacidades de desulfuración son muy altas y equiparables al efecto producido por un lecho de dolomita (con la única excepción de BA a 700ºC). El gran aumento de eficiencia con respecto al char parece estar favorecido, por tanto, por la eliminación del contenido en carbono de los mismos durante la oxidación, quedando así expuestas mayores cantidades de compuestos metálicos activos en la superficie. En cambio, las cenizas de fango presentan resultados de orden similar a los obtenidos con char del mismo material, hecho que puede atribuirse a la similitud entre ambos materiales en cuanto al contenido en cenizas (el char de fango contenía mucho menos C que los producidos a partir de lignito y hulla). La evolución temporal de la eficiencia de desulfuración de las cenizas LA, BA y SSA, junto con la dolomita, se muestra en la Figura 12. En dicha figura puede verse cómo las cenizas de lignito y hulla mantienen su actividad de desulfuración, eliminando la práctica totalidad de H2S durante todo el tiempo de experimento (120 minutos), excepto en el caso de las cenizas de hulla a 700 ºC. Por su parte, las superficies específicas de char y ceniza, tanto antes como después de los experimentos de desulfuración, se muestran en la Tabla 12. 46 Figura 12. Eficiencia de desulfuración (% de S eliminado del gas) de las cenizas y de la dolomita. (a) BA; (b) LA; (c) SSA; (d) dolomita calcinada.  700 ºC;  800 ºC;  900 ºC. Tabla 12. Superficies BET de los materiales empleados. Superficie BET (m2/g) Hulla Lignito Fango Dolomita Char 244.7 173.8 59.1 - Cenizas 5.0 13.0 5.8 18.8 Cenizas tras desulfuración 700ºC 4.9 18.5 - 2.3 800ºC 5.0 10.6 - 8.0 900ºC 3.4 4.5 - 8.3 47 Como puede observarse, una mayor superficie específica no produce un aumento de la eficiencia de desulfuración: las superficies específicas del char de cada material son entre 10 y 50 veces superiores a las de sus cenizas, pero las eficiencias de desulfuración son mucho mayores para estas últimas (excepto, como se ha comentado, en el caso del fango). Además, se determinó la variación de superficie específica de las cenizas de lignito y hulla tras la desulfuración, así como la de la dolomita. Estos datos también se muestran en la Tabla 12 y evidencian cambios poco significativos en cuanto a la estructura porosa de los materiales empleados. Por tanto, a alta temperatura, la superficie BET ejerce una influencia despreciable en la capacidad de desulfuración. Este resultado contrasta con los encontrados en bibliografía empleando materiales carbonosos para desulfuración a bajas temperaturas [42]. Con respecto a este Apartado de la presente Tesis Doctoral, puede concluirse que de los materiales empleados en el estudio de desulfuración, los chares presentan una modesta capacidad de desulfuración, siendo el char de fangos el único capaz de mantener cierta actividad de eliminación de H2S a lo largo del tiempo. Al utilizar cenizas procedentes de lignito y hulla, la eliminación de H2S se ve ampliamente mejorada (reducción de la concentración de sulfuro de hidrógeno de 0.50 a 0.05 % vol. para las cenizas de lignito en todo el rango de temperaturas investigado y para las cenizas de hulla a 800 y 900ºC), y es equiparable a la de la dolomita calcinada, el material de referencia de este estudio. Por su parte, las cenizas de fango sólo muestran una ligera mejora con respecto al char del mismo material. 48 2.4.3. Mejora catalítica del gas usando un catalizador comercial en lecho fijo El objetivo de esta parte de la tesis consistió en el estudio de un sistema de limpieza catalítica del gas de co-gasificación mediante dos lechos fijos consecutivos; uno compuesto por dolomita, y otro por un catalizador comercial de Ni soportado sobre αAl2O3. Como se ha comentado anteriormente, la inclusión de una etapa de desulfuración previa al lecho catalítico de Ni resulta imprescindible, puesto que la presencia de H2S produce la rápida desactivación del mismo por envenenamiento. Es por ello que se optó por emplear dolomita como agente desulfurante de efectividad ampliamente conocida, si bien igualmente podría haberse utilizado alguno de los agentes desulfurantes empleados en la sección 2.4.2; en concreto, cenizas de hulla o lignito. La dolomita presenta además un efecto catalítico de craqueo de alquitranes, de acuerdo con la bibliografía existente [13,43]. Por otra parte, el uso de dos lechos diferenciados para ambos tratamientos permite el control independiente de temperaturas y velocidades espaciales (GHSV, gas hourly space velocity), pudiendo por tanto optimizar estos parámetros para ambos procesos de limpieza del gas. Adicionalmente, se realizaron una serie de experimentos previos incluyendo dolomita dentro del propio lecho fluidizado (sustitución del 25% del lecho de arena por dolomita); sin embargo, esta opción fue desechada debido a la presencia de cantidades muy elevadas de este material en los sistemas de colección de partículas y alquitranes [44]. En los experimentos llevados a cabo en esta etapa, se emplearon cantidades de catalizador suficientes para el estudio de la composición de gases una vez alcanzado el estado estacionario en el sistema, por lo que no se incluyó un análisis detallado de los procesos de desactivación y/o regeneración. Los resultados referentes a la composición de los gases obtenidos se muestran en la Figura 13. Cada experimento se denota de la forma: X-T, siendo X el tipo de lecho 49 empleado (B: blanco, D: dolomita; D-Ni: dolomita + Ni), y T la temperatura de operación del mismo (800, 850 y900ºC). B800 B850 B900 DD-Ni800 D-Ni850 D-Ni900 0.0 0.1 0.2 0.3 0.4 C 2 H n , H 2 S (% vol.) B800 B850 B900 DD-Ni800 D-Ni850 D-Ni900 0 4 8 12 16 H 2 , CO, CO 2 , CH 4 (% vol.) Figura 13. Composición media de los gases obtenidos tras cada lecho ( H2,  CO, ▼ CO2, □ CH4,  C2Hn,  H2S). Con respecto a los experimentos de referencia, la presencia de un lecho de dolomita produce un descenso de la concentración de CO2 y un ligero aumento de la de H2, CO y CH4. Por su parte, los hidrocarburos ligeros (C2Hn) desaparecen casi por completo, lo que puede atribuirse a reacciones de craqueo o reformado promovidas por la dolomita sobre estos hidrocarburos y los alquitranes [45,46]. Por su parte el conocido efecto de desulfuración de la dolomita viene dado por su reacción con el H2S para formar CaS: 50 CaO + H2S → CaS + H2O Ec. 4 Si a continuación se añade un lecho de catalizador de Ni, se observa un incremento en la concentración de H2 y CO y un descenso en la de CO2. Además, el contenido en metano del gas desciende abruptamente. Todas estas observaciones pueden atribuirse al efecto de las reacciones de reformado catalizadas por el níquel [47]. Además de los efectos positivos en cuanto a composición promedio del gas, conviene conocer la evolución temporal de esta composición a lo largo de cada experimento, puesto que sus variaciones pueden indicar posibles cambios de actividad catalítica de los materiales de ambos lechos. Esta evolución temporal se muestra (para el experimento D-Ni800) en la Figura 14. 010 20 30 40 50 60 70 80 0.00 0.01 0.02 0.03 0.04 0.05 C2Hn, H2S (% vol. ) tiempo (min) 010 20 30 40 50 60 70 80 0.0 0.1 0.2 0.3 10 12 14 16 18 H2, CO, CH4, CO2 (% vol.) Figura 14. Evolución de la composición del gas en base seca ( H2, CO, CH4, ▼ CO2,  C2Hn,  H2S). Experimento D-Ni800. 51 Como puede apreciarse, la variación temporal de la composición de los principales gases producidos (CO, CO2 y H2) es poco significativa, pero el metano aumenta de forma constante a lo largo del experimento. Por su parte las concentraciones de H2S y C2Hn permanecen a valores muy bajos y sufren poca variación durante el experimento. La detección de H2S puede indicar cierta desactivación adicional del catalizador de Ni por envenenamiento. La caracterización detallada del catalizador usado en el lecho proporciona resultados que respaldan estas suposiciones [48]. El mencionado aumento de CH4 representa un indicio de posible desactivación del catalizador de Ni [49,50], por lo que la evolución de este gas se estudió en detalle. Así, la Figura 15 muestra la evolución temporal de CH4 en todos los experimentos realizados. A 800ºC, la producción de metano en el experimento blanco y en el correspondiente al lecho único de dolomita son muy similares, lo que sugiere que a esta temperatura la actividad de la dolomita con respecto al craqueo de este gas es aún baja. 010 20 30 40 50 60 70 80 90 0.000 0.005 0.010 0.015 0.020 0.025 0.030 0.035 0.040 CH 4 (g/min) t (min) a) Figura 15. Evolución temporal del CH4 ( blanco,  D-Ni,  D 800 ºC), a) 800ºC, b) 850ºC, c) 900ºC. 52 010 20 30 40 50 60 70 80 90 0.000 0.005 0.010 0.015 0.020 0.025 0.030 0.035 0.040 b) CH4 (g/min) t (min) 010 20 30 40 50 60 70 80 90 0.000 0.005 0.010 0.015 0.020 0.025 0.030 0.035 0.040 c) CH4 (g/min) t (min) Figura 15 (Continuación). Evolución temporal del CH4 ( blanco,  D-Ni,  D 800 ºC), a) 800ºC, b) 850ºC, c) 900ºC. En cambio, tras el paso por el lecho de Ni se produce la práctica desaparición de todo el metano durante los primeros minutos de cada experimento, lo que indica la mayor actividad del catalizador hacia el craqueo de este gas. A temperaturas mayores (800 y 900ºC), este incremento de CH4 se produce más rápidamente, lo que sugiere una mayor velocidad de desactivación. 53 Otro aspecto importante a investigar, relacionado con la composición del gas de gasificación, es el poder calorífico inferior (PCI). En la Figura 16 se representan los valores de PCI (kJ/m3N) obtenidos a partir de la composición media de los gases en cada experimento. En los experimentos en blanco se obtienen valores aproximados de 2000 kJ/m3N a 800ºC, y 2400 kJ/m3N a 850 y 900ºC (diferencias que podrían atribuirse a cierto alcance de las reacciones de craqueo, formando gases combustibles, a las temperaturas más altas del intervalo estudiado). Tras el lecho de dolomita, el craqueo adicional de los productos de gasificación produce un incremento del poder calorífico hasta unos 2800 kJ/m3N, mientras que el lecho de catalizador de níquel sólo produce un aumento significativo a 800ºC. Este hecho puede relacionarse con las mayores concentraciones de H2 y CO encontradas a esta temperatura, como se mostraba en la Figura 13. Temperaturas más altas en el lecho de níquel no producen mejoras apreciables del poder calorífico, en comparación con el lecho de dolomita. B800 B850 B900 DD-Ni800 D-Ni850 D-Ni900 1800 2000 2200 2400 2600 2800 3000 3200 3400 3600 PCI (kJ/m 3 N) Experimento Figura 16. PCI del gas generado. 54 2.4.4. Mejora catalítica del gas usando un reactor de lecho fluidizado de dos zonas Este estudio se usó catalizador preparado en el laboratorio, y no catalizador comercial como en el estudio de limpieza de gas en lecho fijo. El catalizador se caracterizó mediante porosimetría, XRD, TPR y SEM. Se comprobó como el Ni se incorpora al catalizador en forma de espinela de níquel, siendo γ-alúmina la única fase cristalina atribuible al soporte. En la Figura 21 se muestra una microfotografía SEM-EDX que muestra que el níquel se encuentra disperso homogéneamente en la superficie, con una relación Ni/Al de 0.28 constante aproximadamente en toda la superficie escaneada. En los ensayos de actividad del catalizador, realizados en lecho fijo, se comprobó como el soporte tiene cierta actividad en el reformado de hidrocarburos salvo para el metano, y que en presencia de H2S, el catalizador se desactiva, dejando de producirse el reformado de metano. Figura 21. Fotografía SEM-EDX del catalizador preparado 61 Como se ha comentado en el Apartado 2.3.2, tras el gasificador, se situó el reactor de lecho fijo con dolomita para eliminar H2S del gas. El uso del lecho de dolomita calcinada permite además, disminuir la cantidad de alquitrán, que disminuye de 15 g/m3N a 0.21 g/m3N. El reformado adicional que se consigue con el catalizador de Ni en el RLFDZ reduce el contenido en alquitrán del gas a límites no detectables. En el laboratorio, realizando análisis con un cromatógrafo de gases con espectrómetro de masas (GC/MS), se comprobó que este contenido es menor que 2 mg/m3N de naftaleno, que es el compuesto mayoritario detectado en el alquitrán producido, por lo que se puede concluir que la limpieza del gas producido es total. Por lo que respecta a la composición del gas, en las Figuras 22 y 23 se muestra la composición media del gas seco para los compuestos más representativos, obtenida en las distintas configuraciones usadas, tanto por el uso del RLFDZ como el efecto de alimentar, en la zona de regeneración distintos porcentajes de oxígeno para llevar a cabo la combustión selectiva del coque depositado en el catalizador. Se puede observar como el uso de dolomita prácticamente no afecta al contenido en H2 del gas, mientras que la concentración de CO y CH4 aumentan ligeramente, mientras que CO2 y C2Hn (la suma de C2H2, C2H4 y C2H6) disminuyen significativamente. Las tendencias observadas pare el CO2 y CO pueden ser debidas al desplazamiento de la reacción de intercambio inversa así como a las reacciones de reformado de alquitrán provocadas por el lecho de dolomita calcinada. Aunque el uso de alúmina en el RLFDZ aumenta H2 y CO, el efecto se ve incrementado por el catalizador de Ni/γ-alúmina. Con catalizador y sin alimentar oxígeno o con únicamente un 1 % de oxígeno en el gas de regeneración se obtiene un máximo en la concentración del gas para estos compuestos. Sin embargo hay que tener en cuenta en estas circunstancias, el sistema no opera de modo estable. 62 Figura 22. Contenido en hidrógeno (a) y monóxido de carbono (b) en el gas sin reformar () y reformado (). Al aumentar el porcentaje de oxígeno en el gas de regeneración, el sistema se comporta de modo más estable, aunque disminuye la concentración de H2 y CO en el gas. Al usar un 10 % de O2 en el gas de regeneración, la concentración de estos gases cae por debajo de la del gas sin reformar, lo que indica que hay un exceso de oxígeno Gasificador Dolomita Alúmina 0% 0% 1% 2% 6% 10% 0 4 8 12 16 (b) CO (% vol. ) Ni/Alúmina % oxígeno 0 1 6 8 10 12 H2 (% vol.) RLFDZ (a) 63 para la combustión del coque depositado en el catalizador, que llega a la zona de reformado y produce la oxidación del gas de gasificación. Figura 23. Contenido en metano (a), dióxido de carbono (b), y C2Hn (c) en el gas sin reformar () y reformado (). 9 10 11 12 13 14 15 19 20 21 (b) CO2 (% vol.) 0.3 0.6 0.9 1.2 1.5 CH4 (% vol.) RLFDZ (a) Gasificador Dolomita Alúmina 0% 0% 1% 2% 6% 10% 0.00 0.02 0.04 0.06 0.08 0.10 1.0 1.1 (c) % oxígeno C2Hn (% vol.) Ni/Alúmina 64 De los porcentajes de oxígeno en el gas de regeneración utilizados, el 2 % es el que ofrece el mejor resultado en cuanto a aumento del PCI del gas, tal y como se aprecia en la Figura 24 (excluido el 0 % de oxígeno, que como se ha comentado no permite la regeneración del catalizador, dado que no se produce combustión del coque). -100 -50 0 50 100 0% Alúmina Variacion del PCI (%) Ni/Alúmina 0% 1% 2% 6% 10% Figura 24. Variación del PCI del gas (%). Como conclusiones más relevantes de este estudio se puede destacar que se ha conseguido una operación de modo estable del sistema experimental completo. En las mejores condiciones probadas, usando un 2 % en volumen de oxígeno en el gas de regeneración del reactor de dos zonas, se consigue un aumento del 37 % en el poder calorífico inferior y una eliminación completa del alquitrán en el gas, menor que el límite de detección mediante GCMS. 65 66 2.4.5. Trabajos futuros Este trabajo se ha centrado en el estudio de la co-gasificación de fango de EDAR y carbón con aire en lecho fluidizado, dando especial importancia al tratamiento de limpieza de gases. Aunque se han implementado con éxito dos estrategias para la limpieza del gas, en lo que a la eliminación de alquitranes se refiere, y antes de pasar el estudio a otra escala, quedan todavía líneas de mejora para continuar la investigación. Por lo que respecta a la limpieza de los gases, aunque el funcionamiento de los catalizadores ha sido satisfactorio, el trabajo debería completarse con: - Un estudio en profundidad de la vida del catalizador, con tiempos de funcionamiento más largos. - Un estudio de desulfuración con otros materiales, incluso en otras condiciones de operación que pudiesen llegar a simular las condiciones de procesos donde se elimina el azufre, como puede ser el proceso Claus. - Ampliar el rango de contaminantes a otros compuestos, bien con azufre como COS que se conoce que se forman en la gasificación y que por falta de medios materiales no se ha abordado en este trabajo, bien de distinta naturaleza, como es el NH3, que también se forma en la gasificación de compuestos con N en su composición elemental. Por otra parte, este estudio se llevó a cabo en un reactor y en una condiciones que, si bien para la biomasa pueden ser adecuadas, cuando se usa carbón se ha obtenido una conversión del mismo baja, con eficacias energéticas y conversión del carbono a gas que se pueden mejorar, por lo que un estudio ampliando la temperatura de la gasificación y/o incluyendo vapor de agua y mezclas del mismo con oxígeno sería de gran interés, aunque para ello habría que diseñar y construir un sistema experimental distinto. 67 68 3. Bibliografía [1] Manyà JJ, Sánchez JL, Ábrego J, Gonzalo A, Arauzo J. Influence of gas residence time and air ratio on the air gasification of dried sewage sludge in a bubbling fluidised bed. Fuel 2006;85(14-15):2027–2033. [2] Manyà JJ, Sánchez JL, Gonzalo A, Arauzo J. Air gasification of dried sewage sludge in a fluidized bed: effect of the operating conditions and in-bed use of alumina. Energy and Fuels 2005;19(2):629–636. [3] Fonts I, Juan A, Gea G, Murillo MB, Sánchez JL. Sewage sludge pyrolysis in fluidized bed, 1: influence of operational conditions on the product distribution. Industrial & Engineering Chemistry Research 2008;47(15):5376–5385. [4] Fonts I, Juan A, Gea G, Murillo MB, Arauzo J. Sewage sludge pyrolysis in a fluidized bed, 2: influence of operating conditions on some physicochemical properties of the liquid product. Industrial & Engineering Chemistry Research 2009;48(4):2179–2187. [5] Ábrego J, Arauzo J, Sánchez JL, Gonzalo A, Cordero T, Rodríguez-Mirasol J. Structural changes of sewage sludge char during fixed-bed pyrolysis. Industrial & Engineering Chemistry Research 2009;48(6):3211–3221. [6] Casajus C, Ábrego J, Marias F, Vaxelaire J, Sánchez JL, Gonzalo A. Product distribution and kinetic scheme for the fixed bed thermal decomposition of sewage sludge. Chemical Engineering Journal 2009;145:412–419. [7] Aznar M. Estudio de la gasificación con aire en lecho fluidizado de lodos procedentes de estaciones depuradoras de aguas residuales urbanas. Universidad de Zaragoza, 2009. [8] Gonzalo A. Procesado termoquímico de fangos procedentes de estaciones depuradoras de aguas residuales. Universidad de Zaragoza, 2004. [9] Ábrego J. Características y aplicación del sólido obtenido mediante pirólisis de fangos de EDAR. Universidad de Zaragoza, 2010. [10] Fonts I. Estudio experimental de la pirólisis en lecho fluidizado de lodos de EDAR. Obtención y propiedades del producto líquido. Universidad de Zaragoza, 2010. [11] Aznar M, Manyà JJ, García G, Sánchez JL, Murillo MB. Influence of freeboard temperature, fluidization velocity and particle size on tar production and 69 composition during the air gasification of sewage sludge. Energy and Fuels 2008;22(4):2840–2850. [12] Fonts I, Cascarosa E, García G, Aznar M, Ábrego J, Arauzo J. Sewage sludge and meat bone meal valorisation by means of pyrolysis and gasification. In: San José MJ, editor. Advances in chemical engineering. Transworld Research Network, 2010. p. 15–32. [13] Sutton D, Kelleher B, Ross JRH. Review of literature on catalysts for biomass gasification. Fuel Processing Technology 2001;73(3):155–173. [14] Maniatis K. Progress in biomass gasification: an overview, vol. 1. Blackwell Science, 2002. [15] Milne TA, Evans RJ, Abatzoglou N. Biomass gasifier “tars”: their nature, formation, and conversion. report NREL/TP-570-25357. 1998. [16] Abad A, Adánez J, García-Labiano F, De Diego LF, Gayán P. Hot coal-gas desulfurization with calcium-based sorbents in a pressurized moving-bed reactor. Energy and Fuels 2004;18(5):1543–1554. [17] Hepola J, Simell P. Sulphur poisoning of nickel-based hot gas cleaning catalysts in synthetic gasification gas: I. effect of different process parameters. Applied Catalysis B: Environmental 1997;14(3-4):287–303. [18] de Diego LF, Adánez J. Factors affecting the H2S reaction with noncalcined limestones and half-calcined dolomites. Energy and Fuels 1999;13(1):146–153. [19] Corella J, Aznar MP, Gil J, Caballero MA. Biomass gasification in fluidized bed: where to locate the dolomite to improve gasification? Energy and Fuels 1999;13(6):1122–1127. [20] Álvarez-Rodríguez R, Clemente-Jul C. Hot gas desulphurisation with dolomite sorbent in coal gasification. Fuel 2008;87(17-18):3513–3521. [21] Pinto F, Lopes H, André RN, Gulyurtlu I, Cabrita I. Effect of catalysts in the quality of syngas and by-products obtained by co-gasification of coal and wastes. 2: heavy metals, sulphur and halogen compounds abatement. Fuel 2008;87(7):1050–1062. [22] Gascón J. Nuevos reactores de lecho fluidizado de dos zonas para valorización de hidrocarburos. Universidad de Zaragoza, 2006. 70 Anexo II. Artículos publicados. 77 78 Influence of feedstock composition in fluidised bed co-gasification of mixtures of lignite, bituminous coal and sewage sludge G. García a , J. Arauzo b , A. Gonzalo b , J.L. Sánchez b , J. Ábrego b, ⇑ a AEMA Servicios, Pol. Ind. El Pilar, C/Fitero 9, E-26540 Alfaro, Spain b Thermo-Chemical Processes Group, Aragón Institute of Engineering Research (I3A), I+D Building, University of Zaragoza, c/Mariano Esquillor s/n, E-50018 Zaragoza, Spain highlights "Design of experiments to study effect of sewage/coal blends on co-gasification. "Co-gasification produces synergistic effects on product composition (tar). "Improved LHV and gas yields (d.a.f. + air) with addition of sewage sludge. article info Article history: Received 1 October 2012 Received in revised form 8 February 2013 Accepted 18 February 2013 Available online 28 February 2013 Keywords: Co-gasification Sewage sludge Coal Experimental design abstract Energy recovery from sewage sludge can be achieved by several thermochemical processes, including its co-processing with other fuels. In this work, co-gasification of mixtures of sewage sludge with two types of coal (bituminous and lignite) was performed in a laboratory-scale fluidised bed reactor. The influence of the feedstock composition on key parameters of gasification—such as gas heating value and yield, cold gas efficiency and tar generation—was determined. Whereas some of these results can be explained as the sum of individual contributions of each feedstock component, some synergistic effects were also identified. Among these, the decrease of tar yield and the increase of H 2 and CO in the gas suggest that co-gasification of sewage sludge with certain types of coal may be energetically advantageous and improve the process performance. Ó2013 Elsevier B.V. All rights reserved. 1. Introduction In the last decades, the increasing challenges on water management and pollution have led to the application of stricter wastewater treatment policies in most countries. As a result, the number of wastewater plants worldwide has increased dramatically, and so has done the amount of sewage sludge generated in these treatment plants [1]. Because of its origin, sewage sludge contains a significant organic fraction with good potential for energy recovery. Several technologies have been extensively investigated and reviewed in the recent literature, including combustion, co-combustion, gasification and pyrolysis [2–4]. While co-combustion of sewage sludge and coal has been subject of several works [5–13], co-gasification has been shortly reported yet. Research carried out at INETI, Portugal (now Laboratório Nacional de Energia e Geologia) is one of the most significant. Lopes and co-workers [14] and Pinto et al. [15] studied the effect of sewage sludge content (0–100% w/w, mixed with bituminous Puertollano coal), gasification temperature (750–900 °C) and equivalent ratio (0–0.3 at 850 °C) on the gas composition and HCl, NH 3 and H 2 S contents from a fluidized bed gasifier. They also analysed the solid residue after gasification. In the second part of this work [16], the authors gasified sewage sludge–straw blends and compared the results with the ones from the previous article. Wang and Xiao [17] simulated a sewage sludge/coal co-gasification power plant using Aspen Plus. Sewage sludge was fed at 80% moisture with the option of a previous step for drying. Results were compared to sludge gasification, coal gasification and coincineration of both materials. It was found that a mixing weight ratio of 50–67% sludge (80% moisture) outperformed co-incineration, coal gasification or sludge gasification power plants. Liu et al. [18] studied the transformation of phosphorus during gasification in a quartz tube reactor. Sewage sludge powder was mixed with a varying amount of coal powder to make the total phosphorus content in the mixture be 0.5, 1.0, 1.5, 2.0, 2.5, and 2.85 wt.%. The temperature range and the gasifying agent were 900–1300 °C and CO 2 respectively. 1385-8947/$ - see front matter Ó2013 Elsevier B.V. All rights reserved. http://dx.doi.org/10.1016/j.cej.2013.02.073 ⇑ Corresponding author. Tel.: +34 976762962; fax: +34 976761879. E-mail address: [email protected] (J. Ábrego). Chemical Engineering Journal 222 (2013) 345–352 Contents lists available at SciVerse ScienceDirect Chemical Engineering Journal journal homepage: www.elsevier.com/locate/cej Reprinted from Chemical Engineering Journal, Vol 22 (2), G. García, J. Arauzo, A. Gonzalo, J.L. Sánchez, J. Ábrego, Influence of feedstock composition in fluidised bed co-gasification of mixtures of lignite, bituminous coal and sewage sludge, 345-352, Copyright (2013), with permission from Elsevier At the Imperial College of London, Reed and co-workers [19] performed a 2 MWth pilot-plant study of air-blown, pressurised sewage sludge co-gasification with coal at 960 °C. The main objective was to study the distribution of trace elements to the output streams and the effect of temperature in the performance of a downstream hot gas filter. In other works by the same research group, sewage sludge gasification was compared to coal gasification, but these two materials were not co-gasified together. For instance, Paterson et al. [20] described the formation of HCN and NH 3 from gasification of sewage sludge and bituminous coal in a laboratory-scale spouted bed gasifier. Finally, Paterson and co-workers [21] also reviewed the design and development of the bench-scale pressurised fluidized bed reactor used in Ref. [20]. Some advantages over coal gasification have been pointed out in the mentioned works [15]. For instance, co-gasification of sewage sludge and coal mixtures was a better option than co-incineration [17], and caused an increase in fuel conversion and heating value [22]. The objective of this study was to investigate the co-gasification of blends of dry sewage sludge with two different types of coal: bituminous coal and lignite. More specifically, we focused on the influence of the blend composition on some key parameters of gasification, such as cold gas efficiency, gas heating value and composition, specific gas yield and tar generation. Experiments were made in a laboratory-scale fluidised bed reactor, according to a previously developed experimental design for mixtures. In this way, the statistical significance of the results could be adequately evaluated. Although out of the scope of this work, the characteristics and behaviour of the solid products would require further research. For instance, the heavy metal content of the solid product (coming mainly from sewage sludge) should be considered. In this regard, Hernández and co-workers [23] performed leaching tests after gasification of similar sewage sludge, and concluded that thermal treatment significantly stabilized heavy metals in the solid product. Also, it should be noted that the mixture of chars produced as a result of co-gasification of sewage sludge and coals could be used as cheap desulfurization agent [24]. 2. Experimental 2.1. Materials Anaerobically digested and dried sewage sludge was received from a Spanish urban wastewater treatment plant. The bituminous coal and lignite were kindly supplied by the Spanish company Endesa. Hereinafter in this work, these materials will be referred to as BC (bituminous coal), LC (lignite coal) and SS (sewage sludge). Table 1 shows the heating values, proximate and ultimate analyses of the raw materials. Table 2 shows the composition of their ashes. These analyses were performed by ICB-CSIC (Instituto de Carboquímica) in Zaragoza. All the feedstock materials were crushed and sieved to particle sizes between 250 and 500 l m. Sand (SiO 2 ) with a mean particle size between 250 and 350 l m was used as bed material. 20 g of sand per 100 g of total solid feed were added to each blend to improve the feeding of the materials into the reactor. Prior to each experiment, each particular feedstock mixture was blended and dried for 24 h at 105 °C. Air was used as gasification agent. In order to avoid fluidization problems, the possibility of ash melting must be assessed, especially considering the high ash content of some of the materials used in this work. The simplified correlation proposed by Seggiani [25] which was obtained for coal and biomass has been used to calculate the initial deformation temperature, obtaining 1340 °C for BC, 1130 for LC and 1000 for SS (with standard deviation of 80 °C from the correlation). At the gasification temperature of 850 °C, no fluidization problem or slag formation was observed. 2.2. Experimental apparatus The laboratory-scale fluidised bed experimental system has been described elsewhere [24,26]. For all the experiments, the bed temperature was 850 °C, the ratio between actual and minimum fluidization velocity in the bed was V f /V mf = 5, and the equivalence ratio was ER = 0.3 (defined as the ratio between actual oxygen used for gasification and stoichiometric oxygen for a complete combustion). Tar and water were trapped in a series of condensers refrigerated by ice, and in a cotton filter. Water content was determined by means of Karl-Fischer titration, and tar was quantified gravimetrically. Gas composition was measured with an Agilent 3000A MicroGC gas chromatograph. The total amount of gasification gas was quantified using a volumetric gas metre. 2.3. Experimental procedure To establish the influence of the feedstock composition on the gasification outcomes, an experimental design based on a simplex lattice mixture was used. This design was augmented with ternary mixtures to determine the response within the experimental region [27]. A figure showing additional information about the experimental design points can be found in Supplementary Material. Replicated experiments with pure raw materials, their binary mixtures and the centroid (1/3 of each raw material) were carried out in order to determine the experimental error as a basis for the analysis of variance (ANOVA) of the results. Design Expert 7.0.0 software has been used in this work for the statistical analysis and data representation. With the proposed experimental design, it is possible to distinguish whether there is any synergystic or Table 1 Heating values, ultimate and proximate analyses of the raw materials. Analysis method BC LC SS Proximate analysis Moisture (wt.%) ISO-579-1981 6.9 19.2 6.5 Ash (wt.%) ISO-1171-1976 14.3 26.3 41.3 Volatile (wt.%) ISO-5623-1974 25.2 26.0 46.9 Fixed C (wt.%) 53.6 28.5 5.4 Ultimate analysis (dry basis) C (wt.%) Carlo Erba 1108 65.3 37.6 27.8 H (wt.%) Carlo Erba 1108 4.1 4.1 4.4 N (wt.%) Carlo Erba 1108 1.8 0.5 4.0 S (wt.%) Carlo Erba 1108 0.7 6.1 0.8 Heating value (dry basis) HHV a (MJ/kg) ASTM-D-3286-96 25.4 14.6 12.0 LHV a (MJ/kg) ASTM-D-3286-96 24.3 13.3 10.9 a HHV: higher heating value; LHV: lower heating value. Table 2 Ash compositions of the feedstock materials. (wt.% of ash) BC LC SS Al 2 O 3 27.97 25.07 21.57 CaO 7.21 3.40 23.14 Fe 2 O 3 2.71 28.04 8.56 K 2 O 0.74 1.34 3.66 MgO 1.33 1.07 5.93 Na 2 O 0.35 0.16 1.37 SiO 2 47.50 37.88 34.51 TiO 2 1.49 0.73 1.25 346 G. García et al. / Chemical Engineering Journal 222 (2013) 345–352 Reprinted from Chemical Engineering Journal, Vol 22 (2), G. García, J. Arauzo, A. Gonzalo, J.L. Sánchez, J. Ábrego, Influence of feedstock composition in fluidised bed co-gasification of mixtures of lignite, bituminous coal and sewage sludge, 345-352, Copyright (2013), with permission from Elsevier antagonistic effect when the different raw materials are mixed. Thus, the effect of the feedstock composition can be empirically modelled with an equation: RV ¼b 1 W 1 þb 2 W 2 þb 3 W 3 þb 12 W 1 W 2 þb 13 W 1 W 3 þb 23 W 2 W 3 þb 123 W 1 W 2 W 3 ð1Þ In Eq. (1), RV is a response variable (any of the measured experimental results), and the equation represents a special cubic model [28], where the b i coefficients represent the linear response of the pure components mass fraction (W i ), the b ij coefficients the quadratic effect due to the synergetic effect of having two components in the mixture and b 123 models the cubic effect due to the three components. Some of the coefficients may not appear in the model for a specific response, as some of them can be statistically not significant. Thus, this work used this statistical methodology to examine the influence of gasifying different mixtures of the three raw materials on: Gas yield, measured as N m 3 of dry gas produced per kg of d.a.f. feedstock, Y gas , and as the weight percentage of gas related to the feedstock, alone or with air, also in a d.a.f. basis, GY solid (%) and GY solid+air (%). Tar generation (measured either as the tar content in the gas, in g/N m 3 (dry basis), or as the tar yield on a dry and ash free (d.a.f.) basis of the feedstock). Average dry gas composition (vol.% of H 2 , CO, CH 4 ,C 2 H n and H 2 S). Lower heating value (LHV) of the dry gas. Extent of gasification, expressed as carbon-to-gas efficiency, ( g C ), which was calculated as the percentage of carbon in the feedstock converted into C-containing permanent gases (CO, CO 2 ,CH 4 ,C 2 H 2 ,C 2 H 4 and C 2 H 6 ). The cold gas energy efficiency for gasification, g energ , calculated as the ratio between the energy content of the product dry gas and the energy content of the feedstock, as shown in following equation: g energ ¼LHV gas Y gas LHV feed 100 ð2Þ 3. Results and discussion Table 3 shows an overview of the experimental results obtained. At a first glance, it appears that variations of the feedstock composition produce a wide range of variation of all these parameters. For instance, and as can be observed in Table 3, tar content in the gas ranged between 5 and 65 g/N m 3 , with the tar yield in a d.a.f. basis of the solid fed between 1.4% and 17.3% and gas yields between 2.4 and 3.1 N m 3 /kg d.a.f. The lower heating value of the product gas ranged within 1700 and 4100 kJ/N m 3 ). A detailed assessment of the most relevant parameters for gasification will be presented in the following sections. Table 4 shows the statistically significant coefficients, and their 95% confidence interval, as determined by the ANOVA analysis of the experimental results. R 2 is the multiple correlation coefficient also calculated from the ANOVA model, and R 2 adj is the correlation coefficient adjusted for the number of parameters in the model relative to the number of points in the experimental design. R 2 is an indicator of how well the model fits the experimental data. It can be seen that for some responses the regression coefficient is not close to 1, due to the experimental variability. Nevertheless, the statistically significant effect of mixing the three raw materials can be still observed and quantified, although a low regression coefficient means that the obtained model cannot be used for an accurate prediction of the response variables. 3.1. Gas yield Results for the gas yield (Y gas ) are shown in Fig. 1. Firstly, the gas yield (measured as the volume of gas per kg of feedstock on d.a.f.) is shown in Fig. 1a. Y gas is maximum for BC alone, and significantly lower for LC and SS. For bituminous coal–sewage sludge mixtures, Pinto and co-workers [15] observed a similar trend; i.e., a decrease in Y gas at increasing amounts of sewage sludge. For this response variable, no interactions were observed between any of the feedstock components. Thus, Y gas follows a linear relationship with each feedstock component fraction, and can be calculated directly from the feedstock composition using Eq. (1) with the corresponding b i coefficients presented in Table 4. Nevertheless, in this case, low values of R 2 and R 2 adj were determined, as shown in Table 4. This poor fitting of the model limits its usefulness for prediction purposes, and is due to the influence of uncontrolled factors that also might mask some kind of interaction between the feedstock components. The gas yield in a d.a.f. solid basis (GY solid ) is shown in Fig. 1b. The use of air as gasification agent causes these weight yields to be greater than 100%. Using this calculation basis, the highest gas yield is obtained from BC alone, due to the fact that higher air flows are needed with this material for a fixed stoichiometric ratio (mostly because of its higher C content). According to the obtained model, the nonlinear terms are not significant, which means that gas formation is not favoured by any feedstock mixture. Again, it should also be noticed that the experimental variability makes the R 2 value rather low. Table 3 Experimental results (1). Tar, gas, energy and C-to-gas yields, gas LHV and average gas composition (vol.%). Mass fraction BC/LC/SS Tar (g/N m 3 ) Tar (%) g C (%) g energ (%) LHV (kJ/N m 3 ) Y gas (N m 3 /kg) GY solid (%) GY solid+air (%) H 2 CO CH 4 C 2 H n H 2 S 1/0/0 41.2/43.8 12.3/12.6 42.1/42.6 23.2/22.6 2293/2201 3.00/3.13 357/352 80.0/81.2 5.73/5.09 9.04/8.90 1.15/1.16 0.20/0.23 - 0.01/0.04 0/1/0 58.4/64.5 15.6/15.6 47.2/46.6 20.2/22.4 1686/2070 2.67/2.42 337/298 83.6/82.8 5.44/7.50 6.01/7.84 0.80/0.60 0.09/0.10 1.00/0.82 0/0/1 38.4/64.2 10.2/17.3 79.7/79.5 47.3/51.7 4112/3949 2.65/2.69 322/328 86.8/86.6 8.24/7.56 8.52/8.14 3.13/3.14 1.75/1.67 0.10/0.11 0.5/0.5/0 41 11.7 52 22.7 2044 2.86 355 84.6 6.23 7.73 0.92 0.11 0.2 0.5/0/0.5 21.9/28.7 6.1/7.7 56.0/55.0 35.0/33.2 3191/3103 2.81/2.69 339/326 85.6/84.8 8.20/8.09 9.31/9.11 1.97/1.87 0.72/0.69 0.04/0.05 0/0.5/0.5 23.0/18.2 6.0/4.9 64.6/66.7 40.8/42.7 3293/3322 2.62/2.71 317/315 87.2/84.8 8.87/9.04 9.26/9.53 2.23/2.00 0.74/0.72 0.41/0.42 0.67/0.17/0.17 28.4 8.3 48.6 24.3 2302 2.9 359 84.7 6.41 8.25 1.13 0.29 0.07 0.17/0.67/0.17 5.2 1.4 52.7 25 2071 2.77 345 85.5 6.55 7.46 0.86 0.2 0.31 0.17/0.17/0.67 38.1 10.2 64.3 40.5 3477 2.69 325 86.5 8.81 9.32 2.15 0.99 0.14 0.41/0.41/0.18 12.5 3.1 43.3 20.5 2090 2.46 306 81.7 5.9 8.07 0.89 0.2 0.11 0.33/0.33/0.33 6.4/12.9 1.6/3.5 50.1/51.3 29.7/30.5 2882/2731 2.51/2.7 305/329 83.6/83.9 8.10/7.55 9.24/8.87 1.54/1.45 0.49/0.47 0.13/0.32 20.5/20.7 5.4/5.5 46.9/46.4 28.4/27.9 2585/2510 2.62/2.66 319/323 81.0/80.6 6.86/6.65 9.77/9.69 1.10/1.01 0.37/0.35 0.14/0.12 G. García et al. /Chemical Engineering Journal 222 (2013) 345–352 347 Reprinted from Chemical Engineering Journal, Vol 22 (2), G. García, J. Arauzo, A. Gonzalo, J.L. Sánchez, J. Ábrego, Influence of feedstock composition in fluidised bed co-gasification of mixtures of lignite, bituminous coal and sewage sludge, 345-352, Copyright (2013), with permission from Elsevier Calculating the gas yield on an air plus solid d.a.f. basis (GY solid+air ) can provide a more intuitive way of presenting these results. Values are smaller than 100%, and a change in the previously observed trend can be seen: as shown in Fig. 2, the highest yield corresponds to sewage sludge alone, owing to its higher volatile and ash content and lower fixed carbon content, shown in Table 1, especially if compared to bituminous coal. 3.2. Tar generation Results for tar generation are shown in Fig. 3.Fig. 3a depicts the subsequent proposed model for tar content of the gas, according to the coefficients of the adjusted model shown in Table 4. Tar is the main undesirable product obtained from gasification; thus, it should be minimised. From the linear coefficients of the model, it can be seen that both lignite and sewage sludge produce more tar than bituminous coal when gasified alone. It is interesting to notice that the three quadratic coefficients (b ij at Table 3) are statistically significant and have negative sign, indicating that any binary or ternary mixture of the materials causes a decrease in tar yield, if compared to gasification of any of the raw materials alone. This antagonistic effect is presented in Fig. 3a and may be due to the different composition of the inorganics of the three feedstock materials (see Table 2). Al, Ca, Fe and Mg, among other metals, are present in different proportions in the ash of BC, LC and SS, and are widely known to have a catalytic effect towards tar abatement [29–37]. It is likely that a specific combination of these metal compounds within the fluidized bed might optimise tar cracking at these fluidization and temperature conditions. Moreover, differences in the nature of volatiles released and their subsequent gas phase reacTable 4 ANOVA regression coefficients for the measured response variables. Response variable b BC b LC b SS b BC–LC b BC–SS b LC–SS b BC–LC–SS R 2 /R 2 adj Tar (g/N m 3 ) 43±8 58±7 54±7 80 ± 40 90 ± 30 150 ± 30 * 0.75/0.64 Tar yield (%) 11 ± 2 13 ± 2 15 ± 2 * 30 ± 10 40 ± 10 * 0.65/0.54 g energ (%) 23 ± 1 21 ± 1 50 ± 1 2±7 * 7±6 26±6 130 ± 40 0.98/0.97 g C (%) 44 ± 1 47 ± 1 80 ± 1 26 ± 9 23 ± 7 11 ± 7 260 ± 50 0.98/0.97 LHV (kJ/N m 3 ) 2100 ± 90 1900 ± 100 4100 ± 100 *** * 0.94/0.93 Y gas (N m 3 /kg daf ) 2.96 ± 0.06 2.55 ± 0.07 2.63 ± 0.07 *** * 0.55/0.49 Gas yield (%) daf solid 356 ± 7 317 ± 8 316 ± 8 *** * 0.49/0.43 Gas yield (%) daf solid+air 82 ± 1 83 ± 1 87 ± 1 *** * 0.40/0.33 H 2 (%) 6.0 ± 0.4 6.2 ± 0.5 8.4 ± 0.5 ** 5±2 * 0.62/0.55 CO (%) 9.0 ± 0.3 6.7 ± 0.3 8.6 ± 0.4 ** 7±2 * 0.69/0.63 CH 4 (%) 1.1 ± 0.1 0.6 ± 0.1 3.2 ± 0.1 * 1.6 ± 0.6 ** 0.93/0.92 C 2 H n 0.17 ± 0.03 0.069 ± 0.04 1.73 ± 0.04 1.1 ± 0.2 0.8 ± 0.2 * 0.99/0.98 H 2 S 0.05 ± 0.04 0.86 ± 0.05 0.06 ± 0.04 1.3 ± 0.3 0.93/0.92 * Not significant. Fig. 1. Influence of feedstock composition on gas yield: (a) N m 3 /kg on d.a.f. basis; (b) % on d.a.f. basis. Fig. 2. Influence of feedstock composition on gas yield, % on (d.a.f. solid + air) basis. 348 G. García et al. / Chemical Engineering Journal 222 (2013) 345–352 Reprinted from Chemical Engineering Journal, Vol 22 (2), G. García, J. Arauzo, A. Gonzalo, J.L. Sánchez, J. Ábrego, Influence of feedstock composition in fluidised bed co-gasification of mixtures of lignite, bituminous coal and sewage sludge, 345-352, Copyright (2013), with permission from Elsevier tions [38] may also cause different compositions of the formed tars, as pointed out by several authors for coal and wood co-gasification [39,40]. These compositional differences of tars might also enhance the catalytic action of some of the metals present in the ashes. Unfortunately, no further analyses of tar composition were done in this study. Tar yield can also be measured as the weight percentage of feedstock that goes to tar on a dry and ash free basis, as shown in Fig. 3b. This measure seems more convenient for comparison, since it represents the fraction of organic matter converted into tar regardless of the amount of gas formed. As can be seen, there is a quadratic effect caused by the mixtures of SS–LC and SS–BC, whereas the LC–BC mixture is merely linear. The observed antagonistic response (excluding the LC–BC linear behaviour) is similar to that previously shown for tar content in the gas (g/N m 3 ), decreasing the tar concentration. 3.3. Average composition of the dry gas The influence of the feedstock composition on the average content (% by volume) of the main combustible gas components (H 2 , CO, CH 4 ,C 2 H n ), as well as H 2 S (an important gaseous pollutant), is presented in this section. The volume percent (on a dry gas basis) of each one of the main combustible gases are shown in Fig. 4. For instance, the average volume% of H 2 in the dry gas is shown in Fig. 4a. In this work, H 2 changed within the range 5–9%. Paterson and co-workers [22] found somewhat higher concentrations of H 2 and CO, probably because or the use of air/steam mixtures, similarly to the results by Pinto et al. [15]. Wang and Xiao [17] also obtained higher concentrations of these two gases as a result of co-gasifying sewage sludge with high moisture content. It can be seen that the mixtures of LC–BC and SS–BC can be described by a linear combination of the H 2 % produced in the gasification of those materials alone, whereas the SS–LC mixtures show a maximum evidencing a synergistic effect. Unlike our results, Pinto et al. [15] observed a decrease in H 2 concentration with a rise in the sludge content of the feedstock, which could be attributed to the addition of steam. CO contents range between 6% and 10%, as shown in Fig. 4b. The same trend than for H 2 is found: a synergistic effect of feeding SS– LC mixtures and linear effects for the rest of binary mixtures. However, methane shows a slightly different behaviour: for this gas there is an antagonistic (although quantitatively weak) effect between BC–SS, as shown in Fig. 4c. Regarding C 2 hydrocarbons, there is also an antagonistic effect caused by the mixture of sewage sludge with any of the two coals, similarly to the previously shown results for tar generation. The antagonistic effects mentioned for CH 4 and C 2 hydrocarbons are statistically significant although not very important, as can be inferred from the slight curvature shown in Fig. 4d and from the low values of the quadratic terms calculated with the ANOVA analysis (Table 4). Both CH 4 and C 2 hydrocarbons increase with increasing amounts of sewage sludge in the feedstock mixture, as previously noted by Pinto et al. [15], who attributed this fact to the rapid release of volatiles that do not crack substantially due to short residence times in the bed. Finally, results for H 2 S are presented in Fig. 5. The antagonistic effect is created by the combination of both coals, but not with sewage sludge. Additionally, and due to the very high sulphur content of lignite, the formation of this gas is maximum for this individual feedstock. In contrast, the H 2 S contents are much lower for BS and SS as individual feedstock. Moreover, both materials produce very similar H 2 S concentrations in the product gas, due to their almost identical S content, as shown in Table 1. 3.4. Lower heating value of the gas As observed from Fig. 6 and from the significant coefficients in Table 4, the lower heating value (LHV, on a dry basis) of the produced gas can be simply calculated by averaging the LHV contribution of each one of the single components in the feedstock, being the quadratic and cubic effects from the composition not significant. The gas produced by gasification of sewage sludge has a noticeably higher LHV than the produced in the gasification of any of the coals tested. This could be related to the increased release of methane and light hydrocarbons with higher sewage sludge contents, as previously mentioned. 3.5. Carbon-to-gas ( g C ) and cold gas energy ( g energ ) efficiencies The carbon-to-gas efficiency, g C , has been calculated from the gas composition and production, as a way to compute the gasification conversion. As shown in Fig. 7a and b, it is clear that the obtained models for carbon-to-gas conversion and cold gas efficiency are almost identical in shape, as could be expected since both parameters measure the conversion towards gas. In both Fig. 3. Influence of feedstock composition on tar yield: (a) g/N m 3 of dry gas; (b) % on d.a.f. basis. G. García et al. /Chemical Engineering Journal 222 (2013) 345–352 349 Reprinted from Chemical Engineering Journal, Vol 22 (2), G. García, J. Arauzo, A. Gonzalo, J.L. Sánchez, J. Ábrego, Influence of feedstock composition in fluidised bed co-gasification of mixtures of lignite, bituminous coal and sewage sludge, 345-352, Copyright (2013), with permission from Elsevier models, it is shown that sewage sludge exhibits the higher conversion towards gas, whereas the two coals have lower conversions. This fact agrees with the work of Pinto et al. [15] and Paterson et al. [22]. In the case of sludge alone, approximately 80% of the original C content can be converted into permanent gases [33] in the conditions used in this work. Comparing to values reported by Paterson et al. [22], carbon-to-gas and cold gas efficiencies are rather low, which can be explained by the use of an air/steam mixtures as gasifying agent in the mentioned work. Fig. 4. Influence of feedstock composition on gas composition, vol.%: (a) H 2 ; (b) CO; (c) CH 4 ; (d) C 2 H n . Fig. 5. Influence of feedstock composition on H 2 S content (vol.%). Fig. 6. Influence of feedstock composition on the lower heating value of gas (dry basis). 350 G. García et al. / Chemical Engineering Journal 222 (2013) 345–352 Reprinted from Chemical Engineering Journal, Vol 22 (2), G. García, J. Arauzo, A. Gonzalo, J.L. Sánchez, J. Ábrego, Influence of feedstock composition in fluidised bed co-gasification of mixtures of lignite, bituminous coal and sewage sludge, 345-352, Copyright (2013), with permission from Elsevier From the regression coefficients shown in Table 4, it can be observed that the highest gas energy efficiencies (up to 50%) are obtained when SS alone or sludge-enriched mixtures are used. This fact could be due to the higher volatile (and lower fixed carbon) content of sludge compared to the coals used. Accordingly, BC and LC show low cold gas efficiencies (and also tar production, as shown in Fig. 3). As a consequence, chars produced from BC and LC still have high organic (mainly carbon) contents. Results from a previous work [24] (carried out in the same experimental facility, with the same raw materials and at similar operating conditions) showed that the carbon content of chars from gasification were 47% for BC, 25% for LC and 14% for SS. Fig. 7b showing the results of the ANOVA model for cold gas efficiency, depicts a rather flat surface. Thus, there are no relevant synergistic effects when using any mixture of the three raw materials, despite the statistical significance of the parameters shown in Table 4. 4. Conclusions In this work, the influence on several gasification parameters of mixing two types of coal (bituminous and lignite) with sewage sludge was studied with the help of a mixture design. Many of the observed effects on the gasification performance can be simply described as a linear combination of the individual contributions by each feedstock component; however, some synergistic or antagonistic effects were also found. For instance, a noteworthy reduction on tar generation was found when mixtures of coal and sewage sludge were gasified. The type of coal had a minor influence on this reduction. This fact might be due to the combined catalytic effect of the inorganics coming from coals and sludge. Regarding gas composition, the main combustible components of the gas (H 2 and CO, in vol.%) were increased with increasing contents of sewage sludge, with a synergistic effect from lignite-sewage sludge mixtures that caused maximum concentrations of both gaseous products. Methane and C 2 hydrocarbons also increased with increasing contents of sewage sludge. The positive effect of adding sewage sludge was also evidenced on the gas lower heating value, carbon-to-gas conversion and gas energy efficiencies. According to these results, co-gasification of coals and sewage sludge might improve the overall process performance at the operational conditions tested in this work. Appendix A. Supplementary material Supplementary data associated with this article can be found, in the online version, at http://dx.doi.org/10.1016/j.cej.2013.02.073. References [1] A. Kelessidis, A.S. 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Sánchez, Use of different residues for high temperature desulphurisation of gasification gas, Chem. Eng. J. 174 (2011) 644–651. [25] M. Seggiani, Empirical correlations of the ash fusion temperatures and temperature of critical viscosity for coal and biomass ashes, Fuel 78 (1999) 1121–1125. [26] G. García, E. Cascarosa, J.L. Sánchez, Analysis of catalytic co-gasification of two types of coal with sewage sludge and olive bagasse, in: Proceedings of the 17th European Biomass Conference and Exhibition, Hamburg, Germany, 2009, pp. 721–724. [27] J.A. Cornell, Experiments with Mixtures: Designs, Models, and the Analysis of Mixture Data, 3rd ed., Wiley, New York, 2002. [28] G.F. Piepel, J.A. Cornell, Mixture experiment approaches: examples, discussion, and recommendations, J. Qual. Technol. 26 (1994) 177–196. [29] S.-P. Hong, J.-I. Dong, S.-K. Yeo, I.-H. Park, M.-S. Chung, D.-I. Kim, et al., Reduction of tar using cheap catalysts during sewage sludge gasification, J. Mater. Cycles Waste Manage. 13 (2011) 186–189. [30] J.N. Kuhn, Z. Zhao, L.G. Felix, R.B. Slimane, C.W. Choi, U.S. Ozkan, Olivine catalysts for methaneand tar-steam reforming, Appl. Catal. B-Environ. 81 (2008) 14–26. [31] S. Rapagnà, K. Gallucci, M. Di Marcello, M. Matt, P.U. Foscolo, Improvement of gas yield from biomass gasification by using Fe/olivine as gasifier bed inventory, Chem. Eng. Trans. 21 (2010) 415–420. [32] M. Azhar Uddin, H. Tsuda, S. Wu, E. Sasaoka, Catalytic decomposition of biomass tars with iron oxide catalysts, Fuel 87 (2008) 451–459. [33] M. Virginie, C. Courson, D. Niznansky, N. Chaoui, A. Kiennemann, Characterization and reactivity in toluene reforming of a Fe/olivine catalyst designed for gas cleanup in biomass gasification, Appl. Catal. B-Environ. 101 (2010) 90–100. [34] L. Di Felice, C. Courson, D. Niznansky, P.U. Foscolo, A. Kiennemann, Biomass gasification with catalytic tar reforming: a model study into activity enhancement of calciumand magnesium-oxide-based catalytic materials by incorporation of Iron, Energy Fuel 24 (2010) 4034–4045. [35] J.M. de Andrés, A. Narros, M.E. Rodríguez, Behaviour of dolomite, olivine and alumina as primary catalysts in air–steam gasification of sewage sludge, Fuel 90 (2011) 521–527. [36] Y. Sun, J. Jiang, E. Kantarelis, J. Xu, L. Li, S. Zhao, et al., Development of a bimetallic dolomite based tar cracking catalyst, Catal. Commun. 20 (2012) 36– 40. [37] J.J. Manyà, J.L. Sánchez, A. Gonzalo, J. Arauzo, Air gasification of dried sewage sludge in a fluidized bed: effect of the operating conditions and in-bed use of alumina, Energy Fuel. 19 (2005) 629–636. [38] F. Pinto, C. Franco, H. Lopes, R.N. André, I. Gulyurtlu, I. Cabrita, et al., Effect of used edible oils in coal fluidised bed gasification, Fuel 84 (2005) 2236– 2247. [39] I. Aigner, C. Pfeifer, H. Hofbauer, Co-gasification of coal and wood in a dual fluidized bed gasifier, Fuel 90 (2011) 2404–2412. [40] K. Kumabe, T. Hanaoka, S. Fujimoto, T. Minowa, K. Sakanishi, Cogasification of woody biomass and coal with air and steam, Fuel 86 (2007) 684–689. 352 G. García et al. / Chemical Engineering Journal 222 (2013) 345–352 Reprinted from Chemical Engineering Journal, Vol 22 (2), G. García, J. Arauzo, A. Gonzalo, J.L. Sánchez, J. Ábrego, Influence of feedstock composition in fluidised bed co-gasification of mixtures of lignite, bituminous coal and sewage sludge, 345-352, Copyright (2013), with permission from Elsevier 650 G. García et al. / Chemical Engineering Journal 174 (2011) 644–651 Table 6 BET surface of dolomite, CA and LA after the desulphurisation tests. BET (m2/g) 700◦C 800◦C 900◦C CA 4.93 5.01 3.43 LA 18.53 10.59 4.46 DO 2.29 8.00 8.31 of Ca, Na, K and Mg in the ash from sewage sludge is higher than in the ash from both types of coal (Table 4), the H2S removal capacity of sewage sludge ash is the lowest of the three materials. These results indicate that the metallic oxides present in the material are not the unique variable that has influence on the capacity of H2S removal. The H2S removal capacities of ashes from bituminous coal and lignite (CA and LA) are higher than those using char from the same materials (CC and LC), which can be related to the fact that ashes do not contain fixed carbon. Regarding the specific surface of the tested materials, shown in Tables 3 and 4, it seems that this parameter does not play a significant role on the desulphurisation runs at this temperature range. The measured specific surface of CA is 5.0m2/g, about 50 times smaller than that of CC (244.7m2/g). For lignite, the specific surface of its ash and char is 173.8m2/g and 13m2/g, respectively. The specific surface of char from sewage sludge (59.1m2/g) is ten times higher than the specific surface of ash from sewage sludge (5.8m2/g). The BET surface area of the best performing materials (dolomite, bituminous coal ash and lignite ash) was also determined after the desulphurisation experiments, as shown in Table 6. Additionally, the XRD diffraction patterns of these materials are shown in Figs. 6–8. 8070605040302010 2θ MgO CaS b) d) c) CaCO3 Ca(OH)2 CaO a) Fig. 6. XRD diffraction pattern of calcined dolomite (a), and after desulphurisation test at 700 (b), 800 (c) and 900◦C (d). CaSO4 CaS Fe3O4 Fe1-xS 8070605040302010 SiO2 2θ b) d) c) a) Fig. 7. XRD diffraction pattern of lignite ash (a), and after desulphurisation test at 700 (b), 800 (c) and 900◦C (d). According to Table 6, surface area of CA does not show a significant change after desulphurisation, only at 900◦C it falls from 5 to about 3m2/g. The initial surface area of LA (13m2/g) increases slightly up to 18m2/g at 700◦C, and decreases again when the desulphurisation temperature increases (4m2/g at 900◦C), a similar value to that of CA at that temperature. Dolomite, which initially had the highest surface area (18m2/g), has a surface area of only 2m2/g after the 700◦C desulphurisation test. This can be related to the fact that partial re-carbonation occurs at this temperature, as observed in the XRD diffraction pattern shown in Fig. 6. After the desulphurisation tests at 800 and 900◦C, the surface area of used dolomite is 8m2/g. CaS CaSO 4 Ca 5 (PO 4 ) 3 ·(OH) MgO 8070605040302010 2 θ SiO 2 Fig. 8. XRD diffraction pattern of bituminouscoalash(a), and after desulphurisation test at 700 (b), 800 (c) and 900◦C (d). Reprinted from Chemical Engineering Journal, 174, G. García, E. Cascarosa, J. Ábrego, A. Gonzalo, J.L. Sánchez, Use of different residues for high temperature desulphurisation of gasification gas, 644– 651, Copyright (2011), with permission from Elsevier G. García et al. / Chemical Engineering Journal 174 (2011) 644–651 651 In another work [22] using catalytic carbonaceous materials, alsoatlowtemperature,itwasfoundthatsurfaceareaisalsoimportant,but,athightemperaturethisdoesnotseem sobesoimportant, as the best performing materials (LA, CA and calcined dolomite) have lower specific surfaces than the materials used in the cited paper, as shown in Table 4. As could be expected, the XRD diffraction pattern of calcined dolomite sample (Fig. 6) reveals CaO y MgO as the predominant phases, with only small peaks adscribed to Ca(OH)2. At 700◦C, the re-carbonation of CaO to CaCO3might be the cause for the low BET surface measured. At 800 and 900◦C, CaCO3is not observed, but a phase of Ca(OH)2can be clearly distinguished in the pattern. Small CaS peaks can be observed at 700 and 900◦C, but not at 800◦C. On the other hand, Mg only appears as MgO in the diffraction patterns, and no other crystalline Mg compounds were detected. The diffraction patterns of lignite ash, shown in Fig. 7, evidence the more complex nature of this material. Previous to the desulphurisation experiments, the main crystalline phases observed in lignite ash are SiO2, Fe2O3(as hematite, pattern not shown in Fig. 7), and CaSO4(as anhydrite). At 700◦C, besides SiO2(which is the only compound appearing in the four lignite ash diffractograms), Fe3O4and CaSO4(anhydrite) are present. No evidence of other S containing compounds is found at this temperature. At 800◦C, the diffraction pattern show iron and calcium sulphides, whereas at 900◦C only CaS can be observed. ThediffractionpatternsofCA are shown in Fig.8.Asobservedfor lignite ash, the only compound present in all these samples is SiO2. In CA before desulphurisation, CaO (pattern not shown in Fig. 8), CaSO4,Fe2O3(patternnotshowninFig. 8) and MgO are also present as crystalline phases. After desulphurisation, CaS and hydroxylapatite (Ca5(PO4)3(OH)) can be found in the patterns of the used CA at 700, 800 and 900◦C. It must be noticed that no crystalline compounds which can indicate that Mg is active for S retention have been observed in any of the samples of dolomite or ashes. As observed by Seredych et al. [23] for desulphurisation at low temperatures, the presence of Fe and Ca compounds in these kinds of materials improves the desulphurisation capacity. As explained in the previous paragraphs, Ca and Fe sulphides have been detected in the XRD patterns of the used samples. 4. Conclusions The desulphurisation of a synthetic gasification gas with a high H2S content was studied with char from bituminous coal (CC), lignite (LA) and sewage sludge (SSC) in the 700–900◦C temperature range. From these materials, the sewage sludge char was the one able to keep its removal efficiency longer, although a significant H2S percentage remained in the gas. When ash from the same material was tested at the same temperatures, H2S removal was dramatically improved in the cases of bituminous coal (CA) and lignite (LA), whereas only a slight improvement was found for sewage sludge (SSA). The H2S removal capacity of ash from the two coals tested is almost as high as the showed by calcined dolomite, the commercial materialtestedasreference.ThesematerialshavealowBETsurface area and the XRD patterns once used revealed the presence of CaS and FeS. Using CA over 700◦C and LA at any of the temperatures tested, allows reducing the inlet H2S concentration of 0.50vol.% to less than 0.05vol.%. References [1] F. Pinto, R.N. André, C. Franco, H. Lopes, I. Gulyurtlu, I. Cabrita, Co-gasification of coal and wastes in a pilot-scale installation. 1: effect of catalysts in syngas treatment to achieve tar abatement, Fuel 88 (2009) 2392–2402. [2] S. Cheah, D.L. Carpenter, K.A. Magrini-Bair, Review of midto high temperature sulphur sorbents for desulphuration of biomassand coal-derived syngas, Energy Fuels 23 (2009) 5291–5307. [3] X. Meng, W. Jong, R. Pal, A.H.M. Verkooijen, In bed and downstream hot gas desulphurization during solid fuel gasification: a review, Fuel Process. Technol. 91 (2010) 964–981. [4] R.Zhang,R.C.Brown,A.Suby,K.Cummer, Catalytic destruction of tarinbiomass derived producer gas, Energy Convers. Manage. 45 (2004) 995–1014. [5] J. Hepola, P. Simell, Sulphur poisoning of nickel-based hot gas cleaning catalysts in synthetic gasification gas. II. Chemisorption of hydrogen sulphide, Appl. Catal. B: Environ. 14 (1997) 305–321. [6] Z. Abu El-Rub, E.A. Bramer, G. Brem, Review of catalysts for tar elimination in biomass gasification processes, Ind. Eng. Chem. Res. 43 (2004) 6911–6919. [7] A. Abad, J. Adánez, F. García-Labiano, P. Gayan, Hot coal-gas desulphurization with calcium based sorbents in a pressurized moving-bed reactor, Energy Fuels 18 (2004) 1543–1554. [8] R. Álvarez-Rodríguez, C. Clemente-Jul, Hot gas desulphurisation with dolomite sorbent in coal gasification, Fuel 87 (2008) 3513–3521. [9] F. Pinto, H. Lopes, R.N. André, I. Gulyurtlu, I. Cabrita, Effect of catalysts in the quality of syngas and by-products obtained by co-gasification of coal and wastes. 2: heavy metals, sulphur and halogen compounds abatement, Fuel 87 (2008) 1050–1062. [10] W.F. Elseviers, H. Verelst, Transition metal oxides for hot gas desulphurisation, Fuel 78 (1999) 601–612. [11] M.A. Caballero, M.P. Aznar, J. Gil, J.A. Martín, E. Francés, J. Corella, Commercial steam reforming catalysts to improve biomass gasification with steam-oxygen mixtures. 1. Hot gas upgrading by the catalytic reactor, Ind. Eng. Chem. Res. 36 (1997) 5227–5239. [12] T.J. Bandosz, Desulphurisation on activated carbons, in: T.J. Bandosz (Ed.), Activated Carbon Surfaces in Environmental Remediation, Elsevier Ltd., Oxford, 2006, pp. 231–291. [13] W. Yuan, T.J. Bandosz, Removal of hydrogen sulphide from biogas on sludgederived adsorbents, Fuel 86 (2007) 2736–2746. [14] M. Seredych, T.J. Bandosz, Sewage sludge as a single precursor for development of composite adsorbents/catalysts, Chem. Eng. J. 128 (2007) 59–67. [15] J.J. Manyà, M. Aznar, J.L. Sánchez, J. Arauzo, M.B. Murillo, Further experiments on sewage sludge air gasification: influence of the non-stationary period on the overall results, Ind. Eng. Chem. Res. 45 (2006) 7313–7320. [16] M. Seredych, T.J. Bandosz, Desulphurisation of digester gas on catalytic carbonaceous adsorbents: complexity of interactions between the surface and components of the gaseous mixture, Ind. Eng. Chem. Res. 45 (2006) 3658–3665. [17] M.He,B.Xiao,S.Liu,Z.Hu,X.Guo,S.Luo,F.Yang,Syngasproductionfrompyrolysis of municipal solid waste (MSW) with dolomite as downstream catalysts, J. Anal. Appl. Pyrolysis 87 (2010) 181–187. [18] G. García, E. Cascarosa, J.L. Sánchez, Analysis of catalytic co-gasification of two types of coal with sewage sludge and olive bagasse, in: Proceedings of the 17th European Biomass Conference and Exhibition, Hamburg, 29 June–3 July, 2009, pp. 721–724, ISBN 978-88-89407-57-3. [19] W. Yuan, T.J. Bandosz, Removal of hydrogen sulfide from biogas on sludgederived adsorbents, Fuel 86 (2007) 2736–2746. [20] A. Ansari, A. Bagreev, T.J. Bandosz, Effect of adsorbent composition on H2S removal on sewage sludge based materials enriched with carbonaceous phase, Carbon 43 (2005) 1039–1048. [21] J. Aˇ ıbrego, J. Arauzo, J.L. Sánchez, A. Gonzalo, T. Cordero, J. Rodríguez-Mirasol, Structural changes of sewage sludge char during fixed-bed pyrolysis, Ind. Eng. Chem. Res. 48 (2009) 3211–3221. [22] A. Bagreev, T.J. Bandosz, On the mechanism of hydrogen sulfide removal from moist air on catalytic carbonaceous adsorbents, Ind. Eng. Chem. Res. 44 (2005) 530–538. [23] M. Seredych, C. Strydrom, T.J. Bandosz, Effect of fly ash addition on the removal of hydrogen sulphide from biogas and air on sewage sludge-based composite adsorbents, Waste Manage. 28 (1998) 1983–1992. Reprinted from Chemical Engineering Journal, 174, G. García, E. Cascarosa, J. Ábrego, A. Gonzalo, J.L. Sánchez, Use of different residues for high temperature desulphurisation of gasification gas, 644– 651, Copyright (2011), with permission from Elsevier Desulfurization and Catalytic Gas Cleaning in Fluidized-Bed Cogasification of Sewage Sludge−Coal Blends G. García, † A. Monzon, ‡ F. Bimbela, § J. L. Sanchez, § and J. Abrego* ,§ † Agua, Energia y Medio Ambiente, Servicios Integrales, S.L. (AEMA), Polígono Industrial El Pilar, C/Fitero 9, E-26540 Alfaro, Spain ‡ Catalysis, Molecular Separations and Reactor Engineering Group, Department of Chemical and Environmental Engineering, and § Thermochemical Processes Group, Aragon Institute of Engineering Research (I3A), University of Zaragoza, I+D Building, C/Mariano Esquillor s/n, E-50018 Zaragoza, Spain ABSTRACT: Energy recovery from digested sewage sludge can be achieved by means of co-gasification with coal in a fluidizedbed system. In this regard, one of the main hurdles in developing a feasible process is the need for gas cleaning, with special emphasis on desulfurization and minimization of the tar content of the product gas. In this work, high-temperature catalytic gas cleaning was investigated by means of two fixed beds placed in series downstream of the gasification system: the first containing dolomite for desulfurization and primary tar cracking and the second containing a nickel-based catalyst for additional gas reforming. The effect of the temperature on the performance of the Ni catalyst bed (800−900 °C) was assessed. The use of dolomite in a secondary bed at 800 °C allowed for a significant reduction in both tar [15−0.21 g/m3standard temperature and pressure (STP)] and H2S (to less than 0.01%) and an increase in the heating value of the gas [lower heating value (LHV) from 2000 to 2800 kJ/m3STP]. The use of the Ni catalyst decreased the tar content of the gas to undetectable levels. The best results were obtained with the Ni-based catalyst at 800 °C, in terms of enhanced LHV (increasing from 2000 to 3300 kJ/m3STP), gas production, which increased from around 2.40 to 2.75 m3STP/kg on a dry and ash-free basis (daf), and energy requirements for the process. However, some evidence of Ni catalyst deactivation was found when operating under these conditions. 1. INTRODUCTION Because of the constant increase in sewage sludge generation, it is widely accepted that alternative pathways for its valorization need to be investigated. 1 Among the alternatives, co-gasification with coal could be a feasible and environmentally sound option for energy recovery. 2,3 One of the crucial steps to implement gasification at an industrial scale is gas cleaning and conditioning, particularly the removal of particulate matter, sulfur compounds, and tars. The problem of tar elimination has traditionally been tackled by selecting appropriate process conditions and gasifier designs, by means of gas filtering and/or scrubbing or via thermal or catalytic tar cracking. Within the latter, the use of inexpensive catalytic bed materials, such as dolomite, 4−6 limestone, 7 and olivine, 8 has been explored, and also various metal-based catalysts have been tested. 9−12 However, one of the main operational problems in catalytic tar cracking is catalyst poisoning by sulfur. 13 Besides their negative effect on catalysts, sulfur compounds are always undesirable in the product gas, regardless of its final use. Because of the high sulfur content of sewage sludge and some coals, significant amounts of sulfur compounds are expected to be present in the product gas after gasification, with H2S being the most abundant. 14 Among the different approaches that can be taken for sulfur abatement 15 is the use of dolomite, a well-known desulfurization agent 16−19 especially suitable for application in fixed-bed reactors. Furthermore, it also shows significant tar cracking activity. 6,20 Finally, nickel-containing catalysts have been widely studied for tar reforming of the product gas from coal or biomass gasification. 10,21,22 In this work, the effectiveness of these two materials for hot gas conditioning has been tested for the specific case of co-gasification of sewage sludge and coal blends. The use of a secondary fixed bed for gas treatment allows for independent control of the operational conditions to be maintained. These conditions do not necessarily have to coincide with the fluidized-bed conditions, in terms of either temperature or gas hourly space velocity (GHSV). Additionally, the excessive carryover of particles can be avoided. In a recent work, the influence of feedstock composition on air gasification was investigated for blends of bituminous coal, lignite, and sewage sludge. 23 It was found that tar generation was minimized when the three materials were mixed and compared to the gasification yield of each one of the materials used individually. Now that the effect of the feedstock composition has been determined, in this work, attention is mainly focused on the quality of the product gas. Thus, the objective of this work is to assess the performance of a sequential fixed-bed system (dolomite and Ni catalyst) for hot gas conditioning of a producer gas from co-gasification of coal and sewage sludge blends. In the experiments, large enough amounts of catalysts were used to determine the product gas composition that can be obtained after reaching steady-state conditions. To attain a complete development of this process, systematic studies of catalyst deactivation and the potential regeneration procedures will be carried out in further research. Received: February 13, 2013 Revised: April 15, 2013 Article pubs.acs.org/EF © XXXX American Chemical Society Adx.doi.org/10.1021/ef400259g |Energy Fuels XXXX, XXX, XXX−XXX Reprinted with permission from Desulfurization and catalytic gas cleaning in fluidized bed co-gasification of sewage sludge-coal blends, G.García, A. Monzon, F. Bimbela, J. L. Sanchez, and J. Abrego, Energy Fuels, DOI: 10.1021/ef400259g , Publication Date (Web): 18 Apr 2013). Copyright (2013) American Chemical Society 2. EXPERIMENTAL SECTION 2.1. Experimental Setup. Figure 1 shows a schematic diagram of the experimental setup. The fluidized-bed gasification system has been described elsewhere. 19 An electrically heated pipe connects the fluidized-bed system to the fixed-bed catalytic reactor. This is maintained at 450 °C to prevent condensation of tar products on its inner wall. The fixed-bed reactor is 990 mm long and has a diameter of 35.5 mm. The gas enters the fixed-bed reactor from the bottom and passes through a distributor plate to the dolomite bed (250 g of fresh dolomite, with around 133 g after calcination). The Ni catalyst bed (135 g) is placed above the dolomite bed by means of a metal basket with a diameter almost identical to that of the inner reactor wall. Quartz wool is tightly placed above the dolomite bed to prevent fluidization and/or solid elutriation. The two independent heating zones in the electrical furnace mean that different operating temperatures can be set for each bed. After exiting the fixed-bed system, the treated gas passes through a condensation system and a filter for tar and water collection. The water content of the collected liquid product is measured by Karl Fischer titration. The weight of the produced tar is calculated by difference with the water. The composition of the permanent gases is determined by an Agilent 3000 series MicroGC system [model G2801A, equipped with two analysis modules, Plot U with a Plot Q pre-column and MolSieve 5A, and thermal conductivity detectors (TCDs)], and the total gas flow rate is measured by a gas meter. Prior to each experiment involving the reforming Ni catalyst, the catalyst was reduced by introducing a flow of 15% H2and 85% N2 through the reactor system [800 °C, 2 h, 2.82 L standard temperature and pressure (STP)/min], with STP at 0 °C and 1 atm, similar to the reduction conditions proposed by Zhang et al. 22 All of the experiments were carried out during a total time of 90 min, a fixed gasification temperature of 850 °C, and a stoichiometric ratio of 0.3 (air and solid flow rates of 2.8 L STP/min and 1.9 g/min, respectively). The temperature for the dolomite fixed bed was fixed at 800 °C with a GHSV of 2.18 h−1, whereas the Ni catalyst operated between 800 and 900 °C with a weight hourly space velocity (WHSV) of 2.14 h−1. Before starting to feed the coal−sewage sludge mixtures to the fluidized-bed reactor, the dolomite was calcined with air for 2 h and 800 °C in the same reactor system. After calcination, the system was purged with nitrogen and the previously mentioned H2−N2 mixture was introduced for catalyst reduction. To discard possible additional thermal effects from the fixed-bed system on the results, blank experiments were performed by passing the gasification products through the secondary reactor without containing any bed material. These experiments were performed at 800, 850, and 900 °C. The experiments will hereafter be referred to as B800/B850/B900, the so-called “blank experiments”; D, the experiment only involving a fixed bed of dolomite working at 800 °C; or D−Ni800/D−Ni850/D− Ni900, the experiments including the fixed bed of dolomite at 800 °C plus the fixed bed of the Ni catalyst working at a temperature of 800, 850, or 900 °C. 2.2. Materials. 2.2.1. Feedstock. A mixture of sewage sludge (which was supplied thermally dried and anaerobically digested by the wastewater treatment plant “Madrid Sur”, from Madrid, Spain), lignite coal (produced in Teruel, Spain), and bituminous coal (imported from South Africa, with both coals supplied by a power plant) was prepared by mixing identical weights of each component. Previously, the three materials were ground and sieved separately to obtain a uniform particle size of 250−500 μm. The composition and lower heating value (LHV) of the feedstock materials are shown in Table 1. Figure 1. Scheme of the experimental laboratory-scale system. Table 1. Analysis of the Feedstock Materials bituminous coal (BC) lignite coal (LC) dried sewage sludge (SS) Ultimate Analysis C (%) 65.29 37.55 27.84 H (%) a 4.09 4.07 4.39 N (%) 1.84 0.46 3.95 S (%) 0.68 6.07 0.75 Proximate Analysis moisture b 6.93 19.23 6.46 ash c 14.33 26.31 41.29 volatile matter d 25.19 26.01 46.89 fixed carbon 53.55 28.45 5.36 LHV (kJ/kg) e 24.32 13.25 10.87 a Includes hydrogen from water. b ISO 579-1981. c ISO 1171-1976. d ISO 5623-1974. e ASTM D3286-96. Energy & Fuels Article dx.doi.org/10.1021/ef400259g |Energy Fuels XXXX, XXX, XXX−XXXB Reprinted with permission from Desulfurization and catalytic gas cleaning in fluidized bed co-gasification of sewage sludge-coal blends, G.García, A. Monzon, F. Bimbela, J. L. Sanchez, and J. Abrego, Energy Fuels, DOI: 10.1021/ef400259g , Publication Date (Web): 18 Apr 2013). Copyright (2013) American Chemical Society Sand with a mean particle size of 273 μm was used as fluidization material. 2.2.2. Catalysts. Two catalytic materials have been used: dolomite for desulfurization and primary tar cracking and a nickel-based catalyst for gas reforming. Carbonated dolomite was supplied by the Spanish company Calcinor, whereas cylindrical monoliths (17 mm outer diameter, 9.5 mm long with six cylindrical holes, with 3 mm inner diameter each) of a commercial Ni-based catalyst supported on calcium aluminate and La-promoted (BASF SG9301) were used for the catalytic reforming. These materials have been characterized by several analytical techniques. Dolomite was characterized by means of optical emission spectrometry with inductively coupled plasma (ICP−OES) and N2 adsorption. With regard to the Ni-based catalyst, previous to the activation step, the Ni catalyst monoliths were characterized by N2 adsorption, X-ray diffraction (XRD), temperature-programmed reduction (TPR), and X-ray photoelectron spectroscopy (XPS). After reaction, some spent catalyst samples were also characterized by XRD and XPS. The catalyst composition was supplied by the manufacturer. The XRD analyses were made in a Bruker D8 Advance Series 2 diffractometer using a copper anode and a wavelength corresponding to that of copper (λ= 1.5418 Å). The measurements were completed in the 2θrange from 3°to 80°, using a scanning rate of 0.017°/s. Temperature-programmed reduction (TPR) analyses were carried out with a Micromeritics PulseChemisorb 2700 analyzer using a TCD. The samples were previously degasified at 110 °CinanAr atmosphere. The samples were reduced using a gas mixture of 10% H2in Ar (50.14 cm3STP/min), starting at room temperature and setting a temperature ramp of 10 °C/min until a final temperature of 1050 °C was attained. Lastly, the XPS analyses of both fresh and spent catalyst samples were performed using a Kratos Axis Ultra DLD spectrometer, with monochromatic Al Kα(hυ= 1486.71 eV) as the Xray source. 3. RESULTS AND DISCUSSION 3.1. Catalyst Characterization. 3.1.1. Dolomite. Table 2 shows dolomite composition, determined by ICP−OES analysis and carried out at the Instituto de Carboqui mica (ICB−CSIC), Zaragoza, Spain. Table 2 also includes data of the porous structure of the dolomite after calcination, as determined by N2adsorption at 77 K using a Micromeritics TriStar II 3000 V6.08A analyzer, taking a value of 0.162 nm2for the cross-sectional area of the N2molecule adsorbed. 3.1.2. Ni-Based Catalyst. Table 3 shows the catalyst composition (as determined by the manufacturer) as well as its porous structure data (determined by N2adsorption, using the same analyzer and in the same conditions as those used for the dolomite measurements). This catalyst has previously been reported for natural gas reforming. 24 The porous structure of the as-received fresh catalyst (before activation) was determined by N2adsorption. The adsorption isotherm (not shown) can be ascribed to the type IV isotherm of the International Union of Pure and Applied Chemistry (IUPAC) classification, showing the usual hysteresis loop between the adsorption and desorption curves typical of mesoporous materials. The surface area determined by the Brunauer−Emmett−Teller (BET) method is relatively low, although in the range of other Ca-containing Ni/α-Al2O3 catalysts found in the literature, both commercial 25 and prepared in-house. 26 Most of the surface area (ca. 92%) is external, according to the tplot obtained using the Harkins and Jura thickness equation. This is logical given the low porosity of the material, as revealed by the BET value obtained. The fraction of micropores is small, and the average pore diameter is around 15 nm, determined from the pore size distribution obtained using the Barrett−Joiner−Halenda (BJH) technique, hence signaling the mesoporous nature of the catalyst. Figure 2 shows the XRD pattern of the unreduced catalyst sample. The material possesses a high degree of crystallinity, showing very high and narrow peaks. The crystalline phases detected correspond to α-Al2O3, NiO (bunsenite), and a mixed calcium and aluminum oxide, CaAl4O7. Some secondary peaks could also be attributed to another mixed calcium and aluminum oxide, hibonite [CaO(Al2O3)6]. These have not Table 2. Composition of Carbonated Dolomite and Porous Structure of Calcined Dolomite composition (carbonated dolomite) compound wt % Al2O30.09 CaO 30.34 Fe2O30.01 MgO 20.63 CO248.93 porous structure (calcined dolomite) BET surface (m2/g) 18.8 micropore area (m2/g) 1.6 micropore volume (cm3/g) 6 ×10−4 mean pore diameter (nm) 20.0 Table 3. Composition and Porous Structure of the Ni Catalyst composition compound wt % α-Al2O370−75 CaO 5−10 La2O31.5 NiO 10−15 porous structure BET surface (m2/g) 16.8 micropore area (m2/g) 1.3 micropore volume (cm3/g) 4.6 ×10−4 average pore diameter (nm) 15.3 Figure 2. XRD pattern of the fresh Ni catalyst. Energy & Fuels Article dx.doi.org/10.1021/ef400259g |Energy Fuels XXXX, XXX, XXX−XXXC Reprinted with permission from Desulfurization and catalytic gas cleaning in fluidized bed co-gasification of sewage sludge-coal blends, G.García, A. Monzon, F. Bimbela, J. L. Sanchez, and J. Abrego, Energy Fuels, DOI: 10.1021/ef400259g , Publication Date (Web): 18 Apr 2013). Copyright (2013) American Chemical Society been included in Figure 2 for the sake of clarity. Standard diffraction patterns of lanthanum crystalline phases [La2O3, pattern ref 05-0602, and La2NiO4, pattern ref 70-0509, from the Joint Committee on Powder Diffraction Standards (JCPDS) database of the International Centre for Diffraction Data] were compared to the sample diffraction pattern. It could be concluded that no La crystalline phases were present, probably because of the low content of La in the catalyst. Figure 3 shows the TPR profile of the as-received commercial Ni-based catalyst sample. The TPR profile shows a major reduction peak at 420 °C with a shoulder at around 540 °C. This major peak can be ascribed to the reduction of the Ni species that interact with the La-promoted calcium aluminate support. Requies et al. 27 correlated the presence of a peak at a reduction temperature of 532 °Cwiththereductionof“fixed NiO”species. Furthermore, similar TPR profiles of Ni/Al2O3catalysts prepared by impregnation can be found in the literature. 26,28,29 A secondary peak appears at a temperature around 680 °C, and an additional small shoulder can be observed at ca. 820 °C. The peak at 680 °C could correspond to the reduction of NiO in intimate contact with the support. 26 A small shoulder at ca. 820 °C was also observed by Medrano et al. 29 in a Ni/Al2O3catalyst modified with Ca with a high Ca/Al molar ratio (0.5). Hou et al. 26 attributed the presence of a peak at around 830 °C to the reduction of a NiAl2O4phase. However, in the present work, the XRD pattern of the catalyst sample (Figure 2) did not reveal any crystalline phase corresponding to NiAl2O4, and the small shoulder observed in Figure 3 at 820 °C indicates a low H2consumption. Therefore, it could be concluded that such a shoulder could correspond to the reduction of highly dispersed non-stoichiometric amorphous nickel aluminate spinels. The formation of highly dispersed Ni spinels is favored by the addition of lanthanum as a modifier of the catalyst support in Ni/Al2O3catalysts. 30 Because the XPS analyses were conducted both to fresh and spent catalyst samples, the results of the XPS analyses are presented in section 3.3.1.2 (Table 4). According to the spectrum of the fresh catalyst sample, when analyzing the energy levels corresponding to Ni 2p3/2, a major peak at a binding energy (BE) equal to 855.4 eV is found, with a secondary peak at 862 eV. This has been attributed to the presence of superficial NiO having a certain degree of interaction with the support, concordant with the XRD and TPR analyses previously presented and the literature. 31,32 3.2. Preliminary Experiments: In-Bed Use of Dolomite. Before carrying out the experiments using the secondary fixedbed reactor, it was considered of interest to check the possibility of using dolomite directly in the fluidized bed. Thus, preliminary experiments were performed with fixed amounts of dolomite, substituting 25% of the initial weight of the bed material (sand). The use of this material within the fluidized bed dramatically decreased the H2S content of the product gas (from 0.13 to 0.01 vol %). Nevertheless, a large amount of dolomite fines were found in the condensable collection system, and changes in the tar production could not be quantified. This problem has been reported elsewhere in Figure 3. TPR profile of the commercial Ni-based catalyst. Figure 4. Average gas composition (■,H 2;●, CO; ▼,CO 2;□,CH 4;△,C 2Hn; and ○,H 2S). Energy & Fuels Article dx.doi.org/10.1021/ef400259g |Energy Fuels XXXX, XXX, XXX−XXXD Reprinted with permission from Desulfurization and catalytic gas cleaning in fluidized bed co-gasification of sewage sludge-coal blends, G.García, A. Monzon, F. Bimbela, J. L. Sanchez, and J. Abrego, Energy Fuels, DOI: 10.1021/ef400259g , Publication Date (Web): 18 Apr 2013). Copyright (2013) American Chemical Society similar systems. 33 As a consequence of this excessive carryover, the use of in-bed dolomite was definitively discarded. 3.3. Use of a Secondary Fixed Bed with Dolomite and Ni Catalyst. 3.3.1. Gas Composition. 3.3.1.1. Average Gas Composition. The average gas composition of the experiments is shown in Figure 4. The observed trends are consistent with previously published results for sewage sludge gasification. 34 The dolomite bed produces a decrease in CO2and a slight increase in H2. Both CO and CH4contents increase after the dolomite bed, whereas C2Hnhydrocarbons almost disappear. This might be related to steam reforming and/or cracking reactions for C2Hn hydrocarbons and some tar compounds. Dolomite is known to have a certain catalytic activity for tar cracking. 35,36 The low values that can be observed in Figure 4 for the H2S concentration can be attributed to the sulfidation reaction with dolomite 37 according to the following reaction: +→+CaO H S CaS H O 22 (1) Using the Ni catalyst bed, an increase in the H2and CO contents of the product gas and a slight decrease in the CO2 concentration is observed at all of the temperatures tested and the CH4content falls dramatically. All of these observations are a consequence of the reforming reactions catalyzed by Ni. 32 3.3.1.2. Evolution of Gas Composition. Data from Figure 4 show an overall positive effect of both fixed beds on the average gas composition and cleaning. Nevertheless, it would be interesting to detect any possible temporal change in the gas composition during the experiments, because this would provide evidence for changes in the catalytic activity of the fixed beds. Figure 5 depicts the evolution of the gas composition for the experiment carried out with a Ni catalyst bed temperature of 900 °C. The experiments carried out at 800 and 850 °C showed a similar gas evolution. It is worth mentioning that the Ni catalytic beds after reaction showed two differentiated zones with different colors. The first two-thirds of the catalytic bed (situated at the reactor entrance) contained pellets with a much darker color than that of the initial calcined precursors, probably because of deposited coke. These pellets will be referred to as the black catalyst samples. However, the other third (closer to the reactor outlet) had a whitish look. These will be referred to as the white catalyst samples. As seen in Figure 5, the main product gases from cogasification (H2, CO, and CO2) show little variation during the total experiment time. On the other hand, CH4increases monotonically. The concentrations of C2hydrocarbons and H2S remain almost constant and at very low values. The increase in CH4, although still below the concentration observed in the absence of the Ni catalyst, may be an indication of catalyst deactivation, as described in other works. 38,39 Figure 6 shows the evolution of the CH4production rate (g/min) during each experiment. It can be observed that, at 800 °C, the CH4production rate in the blank experiment and after passing through the dolomite bed is roughly the same in both cases, suggesting that, at this temperature, the activity of dolomite toward methane cracking is low. As expected, the Ni catalyst is more active for methane cracking, and as commented before, the CH4increases over time at the three temperatures. It can also be observed, more clearly at 800 and 950 °C, that the higher the temperature, the quicker the increase in CH4. This could be related to a faster deactivation of the catalyst occurring at higher temperatures. This is correlated with the gas production and heating value observed, which is discussed subsequently. With regard to the evolution of H2S with time, negligible amounts of H2S were detected in the gas composition throughout. This might indicate additional deactivation of the Ni catalyst by sulfur poisoning. To explain the tendencies observed, some characterization of the spent black and white Ni catalyst samples was carried out. Both fractions were characterized separately by means of XRD and XPS analyses. Figure 7 shows the XRD pattern of both samples (a) black and (b) white. The differences found in the crystalline phases corresponding to both samples are evident, even though both present high and narrow peaks, indicating a high degree of crystallinity. The crystalline phase corresponding to α-Al2O3(corundum) could be identified in both cases. Figure 5. Evolution of dry gas composition (■,H 2;●, CO; ▼,CH 4;▼,CO 2;⧫,C 2Hn; and ★,H 2S). Experiment D−Ni800. Energy & Fuels Article dx.doi.org/10.1021/ef400259g |Energy Fuels XXXX, XXX, XXX−XXXE Reprinted with permission from Desulfurization and catalytic gas cleaning in fluidized bed co-gasification of sewage sludge-coal blends, G.García, A. Monzon, F. Bimbela, J. L. Sanchez, and J. Abrego, Energy Fuels, DOI: 10.1021/ef400259g , Publication Date (Web): 18 Apr 2013). Copyright (2013) American Chemical Society However, the rest of the crystalline phases do not coincide. The most significant difference is found for metallic Ni (Ni0), which is present in the black sample but not in the white. In addition, an additional crystalline phase of a mixed calcium aluminum oxide (CaAl4O7) was only identified in the black sample, whereas such a phase could not be detected in the white sample. Some secondary diffraction peaks corresponding to bunsenite (NiO) and hibonite [CaO(Al2O3)6] could explain some of the small diffraction peaks detected in the white sample, at 2θ= 62.9°and 79.5°(NiO) and at 2θ= 20.1°, 32.7°, 40.7°, 40.9°,45.1°,58.9°,and60.0°[CaO(Al2O3)6], respectively. However, the low intensities of the peaks corresponding to NiO and [CaO(Al2O3)6] indicate a low degree of crystallinity for these phases in the sample. Table 4 and Figure 8 show the results of the XPS analyses carried out on the black sample (Figure 8a) and the white sample (Figure 8b). For the sake of comparison, the XPS analysis of the fresh catalyst (already discussed in section 2) is also presented in Table 4 and Figure 8c. The region of interest for BE between 850 and 885 eV was analyzed, which can help in determining differences in the electronic state of surface Ni by analyzing the Ni 2p3/2 and 2p1/2 levels and also the characteristic satellite peaks. Deconvolution curves (not shown) of the overall experimental peaks were generated by applying a Gaussian/Laurentian mixed model of variable proportion. This procedure can help identify hidden secondary peaks that might explain some shoulders and non-Gaussian peak shapes observed in the curves. These curves have been fitted and compared to the overall experimental curves. Again, the differences between the black and white samples are evident. The major peak observed for the black sample at a BE around 855 eV presents a shoulder. Hence, the observed peak can actually be considered as the sum of two peaks (see Figure 8a): one peak with its maximum at 855.7 eV and a second peak with its maximum at 853.1 eV. The latter can be attributed to reduced nickel, in agreement with the XRD analysis (Figure 6), whereas the peak at 855.7 eV, along with the satellite peak found at 862.1 eV, could correspond to stoichiometric NiO having a certain interaction with the support, consistent with the TPR analyses previously presented for the fresh catalyst sample. This is also consistent with the observations made by Salagre et al., 31 who proposed that the shift from the BE corresponding to pure NiO (855.0 eV) toward 855.5 eV indicated a weak interaction with the support. In contrast, the white sample presents a major peak with a maximum at 855.3 eV and a very small peak with a maximum at 851.7 eV. The latter was not detected in the XPS analysis of the black sample. The major peak at 855.3 eV, along with the satellite peak found at 861.9 eV, could again correspond to the presence of a surface NiO phase, concordant with the crystalline NiO phase revealed by the XRD analysis (Figure 6b), even though the slight decrease in the BE values found in the white sample indicate a weaker interaction with the support. On the other hand, the small peak at 851.7 eV could correspond to segregated surface Ni. 40 The peak size might indicate that only a small amount of surface Ni is present, in agreement with the literature, 31 which could explain the lack of detection of a crystalline Ni phase in the XRD analysis. It is also worth noting that the presence of surface S (BE between 160 and 175 eV, not shown) was only evidenced in the white sample. The preferential adsorption of sulfur in ceramic supports has also been previously reported in the literature. 41 The crystalline phase corresponding to NiO was not detected in the XRD analysis of the black sample, meaning that the stoichiometric NiO may be affected by some “decorating”effect caused by dispersed species from the support, mainly Al3+ and La3+ ions. 30 Other authors have reported the existence of difficult to reduce NiO, especially if the particles are small and well-dispersed. 28 The intimate contact of a stoichiometric NiO phase “decorated”with the alumina surface could lead to a moderate interaction that might distort the crystalline structure of the stoichiometric NiO, also resulting in varying degrees of Figure 6. CH4evolution (■, blank experiment; ●,D−Ni; and ★,D 800 °C) at (a) 800 °C, (b) 850 °C, and (c) 950 °C. Energy & Fuels Article dx.doi.org/10.1021/ef400259g |Energy Fuels XXXX, XXX, XXX−XXXF Reprinted with permission from Desulfurization and catalytic gas cleaning in fluidized bed co-gasification of sewage sludge-coal blends, G.García, A. Monzon, F. Bimbela, J. L. Sanchez, and J. Abrego, Energy Fuels, DOI: 10.1021/ef400259g , Publication Date (Web): 18 Apr 2013). Copyright (2013) American Chemical Society reducibility of the NiO phase as observed in the TPR. It might also hinder the reducibility of part of the NiO present in the fresh sample. 3.3.2. Gas LHV. The LHVs of the gases are shown in Figure 9. With regard to the blank experiments, values of around 2400 kJ/m3STP were found for B850 and B900, whereas run B800 produced a gas with a LHV of around 2000 kJ/m3STP. This difference can be attributed to some extent to tar cracking at the highest temperatures of the interval and the subsequent formation of gaseous products. In fact, H2, CO, and CH4 concentrations were found to be slightly higher at B850 and B900, as seen in Figure 4. The additional catalytic cracking in the dolomite bed produces an increase in the LHV of the gas to 2800 kJ/m3STP. Afterward, the Ni catalyst bed produces a remarkable LHV increase at 800 °C (3300 kJ/m3STP), but higher Figure 7. XRD analyses of the spent catalyst samples: (a) black and (b) white. Table 4. XPS Analyses of the Ni Catalyst Samples sample primary peaks BE values (eV) secondary peaks BE values (eV) identified phase Ni/Al atomic ratio black 853.1 Ni00.09 855.7 862.1 NiO white 851.7 Ni00.10 855.3 861.9 NiO fresh 855.4 862.0 NiO 0.14 Figure 8. XPS analyses of (a) black, (b) white spent Ni catalyst, and (c) fresh Ni catalyst samples. Energy & Fuels Article dx.doi.org/10.1021/ef400259g |Energy Fuels XXXX, XXX, XXX−XXXG Reprinted with permission from Desulfurization and catalytic gas cleaning in fluidized bed co-gasification of sewage sludge-coal blends, G.García, A. Monzon, F. Bimbela, J. L. Sanchez, and J. Abrego, Energy Fuels, DOI: 10.1021/ef400259g , Publication Date (Web): 18 Apr 2013). Copyright (2013) American Chemical Society temperatures do not cause significant improvements compared to the gas exiting the dolomite bed. This can be related to the higher H2and CO concentrations found at 800 °C, as seen in Figure 4. 3.3.3. Gas Yield. Some of the previously observed changes in the gas composition can also be correlated to the gas yield, shown in Figure 10. The gas yield increase in the blank experiments at higher temperatures can be explained as being a consequence of the cracking of some tar compounds. The dolomite bed produced a descent in the gas yield that might be caused by partial carbonation of the dolomite (consistent with the observed increase in the LHV) and the removal of H2S. With regard to the catalytic runs, the greatest gas yield was obtained at the lowest reaction temperature studied (800 °C), whereas between 850 and 900 °C, no statistically significant differences were found. The higher gas yield obtained at 800 °C compared to those obtained at higher temperatures might indicate a change in the deactivation rate, which is dependent upon the difference in the formation and gasification rates of carbon/coke precursors. 39 Figure 11 shows the changes in individual gas yields expressed in grams of gas per kilograms of feedstock mixture [dry and ash-free basis (daf)]. Because air is used as a gasifying agent, the weight of gas produced is higher than the weight of organic solid fed. The most relevant changes induced by dolomite on the gas yield include a sharp descent in C2Hnand H2S. Also, lower amounts of CO2are found, probably as a result of the previously mentioned partial carbonation of dolomite. Afterward, the Ni catalyst bed produces high amounts of H2and CO (especially at 800 °C) and a substantial drop in CH4. This could be attributed to the catalyzed steam reforming reaction of CH4, according to eq 2. +↔+CH H O CO 3H 42 2 (2) On the other hand, the differences found in the individual gas yields in the catalytic runs with different reaction temperatures may be explained by the change in the total gas yield, which has previously been discussed. 3.3.4. Gasification Efficiencies and Energy Balance. Two important parameters for gasification systems have been calculated, namely, cold gas efficiency and carbon−gas efficiency. Their values are shown in Figure 12. Cold gas efficiency (ηgas) serves as a measurement of the effectiveness of the gasification process for chemical energy conversion, according to eq 3 η=× Y (%) LHV LHV 100 gas gas gas gas (3) where LHVgas and LHVss are the lower heating values of the product gases and sewage sludge, respectively, and Ygas is the previously calculated gas yield. The carbon−gas efficiency measures the effectiveness of C conversion within the gasification process and can be calculated by eq 4. η=×(%) grams of carbon in gas grams of carbon fed to the gasifier 100 C(4) Cold gas efficiencies increase with the use of dolomite and Ni catalyst beds. In the case of the dolomite bed, partial carbonation would lead to removal of CO2and an increase in the LHV of the gases, as previously observed. This would counteract the previously observed decrease in Ygas and produce a net increase in the cold gas efficiency. With regard to the Ni catalyst bed, the most significant increase takes place at 800 °C, as a result of increases in both the LHV of the gases and Ygas. The cold gas efficiency provides an estimate of energy in the feed stream, which is converted into fuel gas, but does not take into account the energy needed for the hot gas conditioning process. Air gasification can be carried out autothermally, because oxidation reactions supply the energy needed for the different steps (drying, thermal decomposition, dry and steam reforming, etc.). Nevertheless, energy requirements for the dolomite and nickel beds should be clarified. To do this, an energy balance has been carried out with the following assumptions: (1) The only reactions taken into account in the solids are sulfidations of dolomite (where only CaO reacts with H2S to form CaS 42 ) and the Ni catalyst. (2) Because of the lack of data, carbon deposition and CaO carbonation have not been taken into account. (3) Only C2H4and naphthalene have been considered for C2Hnand tar, because they are the major compounds. As seen in Table 5, the enthalpy needed in the dolomite bed (546.8 kJ/kg daf) is 8.1% of the final gas LHV. This value decreases to 6.0% if a Ni bed is used, because of the additional increase in LHV after this treatment. Combining both cleaning steps at 800 °C (dolomite + Ni) would consume 9.0% of the Figure 9. LHV of the product gas. Figure 10. Gas yield. Energy & Fuels Article dx.doi.org/10.1021/ef400259g |Energy Fuels XXXX, XXX, XXX−XXXH Reprinted with permission from Desulfurization and catalytic gas cleaning in fluidized bed co-gasification of sewage sludge-coal blends, G.García, A. Monzon, F. Bimbela, J. L. Sanchez, and J. Abrego, Energy Fuels, DOI: 10.1021/ef400259g , Publication Date (Web): 18 Apr 2013). Copyright (2013) American Chemical Society In the following sections, the different parts of the experimental setup are explained in detail. 2.3.1. Cogasification Section. The cogasification fluidized bed reactor is made of refractory steel (AISI 310), with an inner diameter of 40 mm in the bed zone and 70 mm in the freeboard. The maximum bed height was kept at 310 mm by means of a lateral pipe, which enabled the continuous removal of ash and char. The feedstock blend was continuously fed into the bed at about 10 mm over the distributor plate through a sloped pipe externally cooled by air. The air flow rate was 2.81 dm3(STP)/min (standard temperature and pressure (STP): 0°C and 1 atm). The feeding rate was set at 2.3 g/min (to obtain a stoichiometric ratio of 30%), but slight variations caused the stoichiometric ratio to range between 28.8 and 32.2% depending on the experiment. Both bed and freeboard temperatures were set at 850 °C by means of an electric oven. The initial bed was composed of 300 g of silica sand, with a particle diameter range of 250−350 μm. The product gas exited the reactor through a cyclone and a hot filter, which were kept at 450 °C. The total experiment took 90 min after starting the feeding. 2.3.2. Desulfurization Section. The desulfurization section is composed of the desulfurization reactor and the first gas cleaning and gas analysis system. Gasification gas, as produced, was led to the fixed bed desulfurization reactor through a steel pipe heated at 450 °C to prevent tar condensation. This reactor is also made of refractory steel, with an inner diameter of 35.5 mm and a total height of 990 mm. Calcined dolomite (135 g) was placed as bed material inside the reactor in order to retain the hydrogen sulfide. An electrical oven kept the bed at 800 °C. 2.3.3. Reforming, Gas Cleaning, and Analysis Sections. After the secondary reactor, the gas is divided into two streams. The highest fraction of the gas (“non-reformed flow”, about 3 dm3(STP)/min) exits the experimental installation through the first condensation and gas cleaning system. The rest of the gas (“reformed flow”, 0.3 dm3(STP)/min) is introduced inside the reforming reactor where a bed of Ni/γ-Al2O3catalyst (42 g) is placed. As explained in the Introduction, the reforming reactor chosen is a TZFBR (Figure 3). It was designed in order to combine in only one vessel the two typical phases of catalytic processes: catalytic reaction and catalyst regeneration. This reactor is made of quartz and has a total height of 510 mm. The regeneration zone is 18 mm in diameter, whereas the reaction zone is 28 mm in diameter. The gas to be reformed enters the reactor from the top through a quartz tube, which directs it to the lower part of the reaction zone. The regeneration gas enters through the bottom of the reactor, as can be observed in Figure 3. Reforming reactions take place in the upper part of the bed, while regeneration reactions (coke combustion) occur in the lower fraction of the bed by means an oxygen flow introduced from the bottom of the reactor. An important parameter for the performance of this reactor is the u/umf ratio (actual gas superficial velocity/minimum fluidization superficial velocity) in the two sections of the bed. 52 In the experiments carried out, in order to achieve the appropriate fluidization conditions, the oxygen stream was fed diluted in nitrogen, with a total flow of 0.3 dm3(STP)/min in the regeneration zone (u/umf = 2.5). Nitrogen was chosen for the sake of experimental simplicity, but it must be borne in mind that the reformed gas is subsequently diluted in nitrogen. This dilution effect has been removed by subtracting the nitrogen flow in order to show the effect of the catalyst by comparing the nonreformed gas composition with the reformed one. In an actual operation, either a TZFBR with different dimensions, which enables the fluidization of the catalyst in the regeneration zone of the TZBFR using non diluted oxygen, or oxygen and steam as regeneration agents should be used. In the latter case, steam may easily be removed as water when cooling the reformed gas. The gas flow rate from the desulfurization section entering the regeneration zone of the TZBFR was approximately 0.3 dm3(STP)/ min, which represents a u/umf = 3.4 (taking into account both the larger diameter of this section and the gas streamflow rate coming from the regeneration zone). Because of the continuous circulation of the catalyst between the two zones, coke deposited over its surface during the catalytic reforming and cracking reactions is burned in the regeneration zone. Before the start of the experiment, the catalyst was reduced in a H2 stream (vol 5% in N2) at 800 °C, in the same way as in the catalyst activity tests. The operational temperature of the TZFBR for the experiments was also set at 800 °C. This temperature was chosen in order to promote CH4reforming and tar cracking. 2.3.4. Double Gas Cleaning and Analysis System. The “nonreformed flow”was cleaned with two ice-cooled condensers and a cotton filter (first condensation system) in order to remove water and tar from the gas. Once cleaned, the total gas volume was measured and its composition analyzed with a micro-gas chromatograph (MicroGC1, Agilent 3000A). The “reformed flow”was cleaned with a second condensation system, and its flow rate was measured. Another microgas chromatograph (MicroGC2, Agilent 3000A) was used to analyze the composition of this “reformed gas flow”. For both gas flows, the nonreformed and the reformed, H2, CO, CO2,CH 4,C 2H2,C 2H4, C2H6, and H2S concentrations were determined. The gas lower heating value (LHV) of both streams was calculated by means of the determined gas compositions. The tar content of the condensed liquid was analyzed by weight difference between the total liquid condensed and the water content determined through Karl Fischer titration (Mettler Toledo V-20). The tar components were analyzed by gas chromatography with simultaneous mass spectrometry and flame ionization detectors (GC/MS-FID) using an Agilent chromatograph model 7890A, in a similar mode as in previous works. 53,54 GC-FID analyses were used to find out the relative proportion of each compound in the sample by calculating the area percentage. The quantification method by FID area percentage considers the response factors of all the compounds to be similar. This method has been widely used by other authors because when the compounds belong to the same families the response factors do not change significantly. 55 Figure 1. SEM-EDX scan of a Ni-γ-Al2O3catalyst sample. Table 3. Composition of Gas Mixtures Used in Fixed Bed Catalytic Tests SG-1 SG-2 H2(vol %) 9.00 10.00 CO (vol %) 15.00 10.00 CO2(vol %) 11.00 15.00 CH4(vol %) 4.00 4.00 C2H2(vol %) 0.75 0.20 C2H4(vol %) 0.75 1.50 C2H6(vol %) 0.75 0.20 H2S (vol %) 0.50 N2(vol %) 58.75 58.60 Energy & Fuels Article dx.doi.org/10.1021/ef400227z |Energy Fuels XXXX, XXX, XXX−XXXC Reprinted with permission from Gas Catalytic Upgrading In A Two Zone Fluidized Bed Reactor Coupled To A Co-Gasification Plant, G. García, E. Campos, I. Fonts, J. L. Sanchez, and J. Herguido, Energy Fuels, DOI: 10.1021/ ef400227z , Publication Date (Web): 18 Apr 2013). Copyright (2013) American Chemical Society The double gas cleaning and gas analysis system enabled both the tar of the reformed and nonreformed flow and also the reformed and the nonreformed gas composition in each experiment to be determined. This double system was aimed to facilitate the comparison and comprehension of the effect of the TZFBR catalytic processing. 2.4. Experimental Planning. Apart from the catalytic activity tests in the fixed bed reactor, the experiments carried out in this work were gathered in two sets: first, the experiments designed to study separately the effect of each part of the installation (cogasification, desulfurization and gas reforming units) and, second, those experiments focused on finding the oxygen flow necessary to regenerate online the catalyst in the reforming unit, achieving a steady state operation. 2.4.1. Experiments Designed to Study Separately Each Part of the Installation. The first experimental set included some experiments to determine the original response of the cogasification system before any amendment was made to it. The desulfurization reactor with dolomite was then added to the installation and experimentally studied. The last experiments in this series included the addition of the TZFBR with an alumina bed, as a blank study to test the catalytic effect of the catalyst support. Furthermore, several experiments were run under the same operating conditions. The results obtained showed that there is practically no difference between them, confirming the good repeatability of the experiments as can be observed in some of the figures shown in the Results section (Figures 5, 7, and 9). 2.4.2. Oxygen Optimization Experiments. The regeneration oxygen flow was optimized in order to achieve, in the TZFBR regeneration zone, the continuous combustion of the coke deposited on the catalyst surface. This will be achieved but oxygen should not reach the reaction zone. If that happened, the combustible components of the gas in that zone would react with the remaining oxygen after the regeneration, leading to a lowering of the gas quality. The relative oxygen percentages tested were 0 vol % (reference), 1 vol %, 2 vol %, 6 vol %, and 10 vol %. The correct oxygen percentage was determined by trial and error. In the two series of experiments (experiments focused on studying each part of the installation and experiments focused on finding out the oxygen flow required in the TZFBR), the response variables studied were tar production and composition, gas composition, and lower heating value (LHV). The operational conditions used in the two series are shown in Table 4 and were mostly scheduled according to the optimum values obtained in previous works. 25,47,48,56−58 2.5. Characterization of the Used Catalyst. The samples of the catalysts used in the different runs were characterized by controlled thermogravimetric oxidation (30−900 °C, 10 °C/min, 30 cm3(STP)/ min air) using a STA 449 F3 Jupiter (Netzsch) and the CO2present in the gas exiting from the TG equipment was analyzed online using a Pfeiffer mass spectrometer model Omnistar Prisma. 3. RESULTS AND DISCUSSION 3.1. Catalyst Activity. As stated previously, catalytic activity tests were carried out in a fixed bed reactor loaded with alumina or Ni/alumina and with one or other of the synthetic gas streams SG-1 or SG-2. In all the runs carried out with the SG-1 gas, i.e., without H2S as a constituent, steady state was approximately attained in about 10 min and continued until the end of the experiment (150 min). However, in the runs carried out with Ni/alumina catalyst and using the SG-2 gas (i.e., with H2S as a constituent), the steady state was never reached during the testing time. This is consistent with the deactivating effect of H2S found by several authors working with Ni-based catalysts or with oxide catalysts in gas reforming processes, 31,38 and even in gas coming from gasification Figure 2. Experimental plant diagram (TC: temperature controller. FC: flow controller). Energy & Fuels Article dx.doi.org/10.1021/ef400227z |Energy Fuels XXXX, XXX, XXX−XXXD Reprinted with permission from Gas Catalytic Upgrading In A Two Zone Fluidized Bed Reactor Coupled To A CoGasification Plant, G. García, E. Campos, I. Fonts, J. L. Sanchez, and J. Herguido, Energy Fuels, DOI: 10.1021/ ef400227z , Publication Date (Web): 18 Apr 2013). Copyright (2013) American Chemical Society upgrading processes 37,39,59 similar to the current one. Figure 4 shows the changes for each one of the gas species (in vol %) when SG-1 or SG-2 are subjected to the activity tests. The outlet gas composition (vol %) considered was the steady state when working with the SG-1 gas, and the average for last 30 min in the case of working with the SG-2 gas. As can be observed in Figure 4, broadly speaking, methane, carbon dioxide, and C2hydrocarbons act as reactants while hydrogen and carbon monoxide are produced in the process. This behavior is mainly due to the dry reforming of both methane (CH4+CO 2↔2CO+2H 2) and C2hydrocarbons (C2Hn+2CO 2↔4CO+n/2H2). However, the extent of these reactions depends on the solid used as catalytic bed and also on the presence of hydrogen sulfide in the synthetic gas stream. Thus, when only catalyst support (i.e., γ-Al2O3) is used, apparently there is no methane reforming but almost complete C2hydrocarbon reforming. Indeed, taking into account the inlet gas composition, the stoichiometry of the C2reforming reaction and the degree of disappearance of C2(87.5%), a carbon dioxide decrease of 35.6% would be expected, which roughly corresponds with what was observed (Figure 4). Conversely, Ni/γ-Al2O3appears to be able to reform both methane and C2hydrocarbons in the gas, which should theoretically cause a 77.3% decrease in CO2, while experimentally an 82% reduction was determined. This resulted in a significant increase in the volumetric percentage of hydrogen and carbon monoxide in the gas mixture. As the synthetic gas used in these runs was the same, SG-1, it can be said that the catalytic activity of alumina in the reforming of methane and the production of carbon monoxide and hydrogen is much smaller than that showed in the presence of Ni. When H2S is present in the gas (SG-2), as would be expected 31,37−39,59 the Ni/γ-Al2O3no longer reforms methane due to the poisoning of the Ni active sites, but it maintains a good activity in C2hydrocarbons reforming. This corroborates the fact that Ni/γ-Al2O3reforms both methane and light hydrocarbons, but if Ni is poisoned by the sulfur then only light hydrocarbons are reformed by the effect of γ-Al2O3due to the selective deactivation mainly affecting the ability of catalytic methane dry reforming. Thus, the removal of H2S is a necessary task in order to carry outgas upgrading with this Ni/γ-Al2O3 catalyst in a stationary mode. 3.2. Cogasification and Desulfurization. Before extracting results from the experimental plant, some previous experiments were carried out in order to know how the gas quality changes when the different units are added to the cogasification plant. These previous experiments comprised a cogasification process with a single fluidized bed and cogasification with a subsequent desulfurization process using a downstream fixed bed of dolomite. The gas produced in the cogasification process had, on average, the composition and heating value shown in Table 5. As far as tar production is concerned, the gas had around 15 g/ m3(STP) when leaving the gasifier. Table 6 shows the main compounds identified, as well as their FID area percentages, which are about 90% in FID area of the total peaks detected. As can be observed, naphthalene is the major compound, with all the compounds detected having a marked aromatic character. Figure 3. Scheme of the TZFBR. Table 4. Operational Conditions for Each Plant Section cogasifier bed temp. 850 °C freeboard temp. 850 °C cogasification air 2.82 dm3(STP)/min coal + DSS feeding rate 2.3 g/min initial sand bed 300 g stoichiometric ratio 30% desulfurization reactor bed temp. 800 °C dolomite bed (prior to calcination) 250 g calcined dolomite bed 135 g calcination conditions 800 °C, air, >2 h TZFBR bed temp. 800 °C catalyst in reaction zone 25 g catalyst in regeneration zone 17 g gas entering reaction zone ≈0.3 dm3(STP)/min O2+N 2regeneration flow 0.3 dm3(STP)/min Figure 4. Gas composition variation (vol %) in catalytic tests in fixed bed reactor. Energy & Fuels Article dx.doi.org/10.1021/ef400227z |Energy Fuels XXXX, XXX, XXX−XXXE Reprinted with permission from Gas Catalytic Upgrading In A Two Zone Fluidized Bed Reactor Coupled To A CoGasification Plant, G. García, E. Campos, I. Fonts, J. L. Sanchez, and J. Herguido, Energy Fuels, DOI: 10.1021/ ef400227z , Publication Date (Web): 18 Apr 2013). Copyright (2013) American Chemical Society On the other hand, the gas as produced in the gasifier contained 1100 ±200 ppm of H2S (Table 5). In order to reduce the presence of this gas, noxious for the Ni catalyst as previously established, the produced gas was fed into the secondary fixed bed reactor with dolomite, at 800 °C. After the secondary reactor with dolomite, in the experiment made exclusively with this guard bed, the H2S concentration in the gas fell to 140 ±10 ppm. The concentration of tar in the gas flow after the dolomite guard bed was 0.21 g/m3(STP). This represents 98.7% tar reduction by the introduction of the dolomite bed. This tar reduction is slightly higher than those obtained in other works 22 (around 90%), although the ratio of kg dolomite/kg biomass (daf)/h used in this work (1.52) was also greater than that utilized by these authors (0.85 at most). The hydrocarbon cracking effect of dolomite was also evidenced in the dramatic decrease of C2hydrocarbons (C2H2+C 2H4+C 2H6) in the gas, which was 0.31 ±0.05 vol % after the cogasifier and diminished to 20 ppm after the dolomite bed. 3.3. Catalytic Gas Upgrading with the TZFBR. As explained in the Materials and Methods section, a slip stream of the product gas exiting the desulfurization reactor was fed into the TZFBR. First, to distinguish between the effect of the support (γ-alumina) and the catalyst (Ni/γ-alumina), an alumina bed was placed in the TZFBR instead of the catalyst, and the reactor was operated as a conventional fluidized bed (no regeneration oxygen flow was used). The operation conditions were those as previously described (Table 4). Subsequently, catalytic experiments were carried out with the prepared Ni/γ-alumina catalyst and a regenerating stream with different oxygen percentages in order to allow burning of the coke deposited on the catalyst, without burning the reformed gas components. Regarding the condensed liquids remaining after the gas passed through the TZFBR, no tar species were detected by GC/MS-FID in any of the experiments performed using this reactor. The effects of using the TZFBR on the gas compositions are discussed and compared with those previously obtained (i.e., after cogasification and desulfurization units). The volumetric percentage of each gas species in the nonreformed gas stream (gas stream that does not pass through the TZFBR) should be similar in all the experiments and also analogous to that determined in the experiment carried out only with the dolomite guard bed. However, as will be observed in Figures 5, 7, and 9, some differences have been found in these concentrations. These alterations could be related to the different stoichiometric ratios, which varied between 28.8 and 32.2%, in the case of the major species (H2, CO, CO2,CH 4) and also to the difficulties in the quantification of species with low concentrations in the case of C2Hnand H2S. Considering H2and CO as representative products in the reforming process (as previously stated in relation to the catalytic activity tests carried out with synthetic gases), their evolution with the different experimental configurations used is shown in Figure 5. In order to make this comparison, the average percentage of each individual gas has been calculated using the last values of the time-on-stream for each run (the last 30 min in 90 min experiments). The confidence interval for Table 5. Main Properties of Cogasification Outlet Stream avg. gas composition (vol %) H27.9 ±0.2 CO 9.1 ±0.1 CO214.5 ±0.3 CH41.2 ±0.1 C2Hn0.31 ±0.05 H2S 0.11 ±0.02 tar amount in gas (g/m3(STP)) 15 LHV (kJ/m3(STP)) 2590 ±120 Table 6. Tar Compounds Identified by GC/MS-FID cmpd FID area (%) naphthalene 58.3 acenaphthylene 7.1 phenanthrene 3.8 benzo[b]thiophene 3.7 benzonitrile 2.0 quinoline 2.0 dibenzofuran 1.4 fluorene 1.4 biphenyl 1.2 indene 1.1 naphthalene, 2-methyl1.0 fluoranthene 1.0 pyrene 0.8 naphthalene-2-carbonitrile 0.6 anthracene 0.5 naphthalene, 1-methyl0.5 dibenzothiophene 0.4 4H-cyclopenta[def]phenanthrene 0.2 1,8-anthracenediamine 0.2 Figure 5. Hydrogen (a) and carbon monoxide (b) content in nonreformed gas (■) and in reformed gas (○) for different reaction configurations, including several oxygen contents (0−10%) in regenerating gas fed to the TZFBR. (Lines for visual help.) Energy & Fuels Article dx.doi.org/10.1021/ef400227z |Energy Fuels XXXX, XXX, XXX−XXXF Reprinted with permission from Gas Catalytic Upgrading In A Two Zone Fluidized Bed Reactor Coupled To A CoGasification Plant, G. García, E. Campos, I. Fonts, J. L. Sanchez, and J. Herguido, Energy Fuels, DOI: 10.1021/ ef400227z , Publication Date (Web): 18 Apr 2013). Copyright (2013) American Chemical Society these average values, taking into account the Student’stdistribution and standard deviation of the data, has also been calculated and is shown in Figures 5, 7, and 9. In these comparisons, there is one value for the experiments done without the TZFBR and two values for those experiments carried out using the TZFBR: nonreformed and reformed gas. The use of dolomite in the secondary reactor does not affect the H2content in the gas, remaining at around 8% (Figure 5a). Nevertheless, the proportions of CO and CH4slightly increase (from 9 to 11% and from 1.1 to 1.4%, respectively), while CO2, C2Hn, and H2S decrease significantly (Figures 5, 7, and 9). The trends found in CO and CO2could be due to the enhancement by the dolomite bed of both the reverse water−gas shift reaction and the tar reforming reactions. Alumina in the TZFBR increases H2and CO, and the use of Ni/γ-alumina catalyst also has a major positive effect on their concentrations. The maximum contents in these products are obtained with this catalyst and without oxygen or with only 1% of oxygen diluted in the regenerating stream entering the TZFBR. However, in these experimental conditions, a nonstable system is found, as is manifested by the prolonged 95% confidence interval bars. If the oxygen percentage in the regenerating gas increases, the system stabilizes (shorter confidence interval bars), although this decreases the CO and H2concentration in the resulting gas. When 10% of oxygen is used, these CO and H2concentrations fall to values lower than those found in the nonreformed gas. This oxygen percentage is therefore too high, and the excess oxygen reaches the upper zone of the TZFBR, allowing oxidation reactions of the fuel compounds in the gasification gas. Therefore, an optimal oxygen percentage must be selected so that both process stability and gas quality improvement are preserved. For the experimental series shown in Figure 5, a percentage of 2% of O2represents this optimal value taking into account that it is the minimum oxygen content in regenerating gas that allows a stable operation in steady state. In fact, analyzing the time evolution of the CO content in the reformed gas, considering it as a representative specie of the gas composition, it can be seen (Figure 6) that 2% is the lowest oxygen percentage leading to a stable behavior. Working with 2% O2in the TZFBR, during the first 40 min the gas composition changes but for the rest of the experiment the gas composition is very stable. Furthermore, comparing the data of both reformed and nonreformed flows, it can be observed (Figure 5) that the catalyst has a positive effect on the gas composition because the percentages of H2and CO of the reformed gas flow are significantly greater than those of the nonreformed flow. Higher oxygen percentages lead to stable states but lower quality gas. Lower oxygen percentages lead to unstable states, even in long-term experiments (e.g., for 0% and 1% oxygen as shown in Figure 6). In summary, 2% of O2was selected as the most suitable percentage to be used in the regeneration agent gas introduced in the TZFBR due to both the lack of repeatability (or stability) among consecutive measurements of CO and the operation in nonsteady state produced with 0 and 1% of O2. On the other hand, percentages of O2higher than 2% lead to the combustion not only of the coke deposited over the catalyst but also of part of the gasification gas entering the TZFBR. The evolution of the CH4,CO 2, and C2Hnconcentrations, the gas components expected to act as reactants (i.e., decreasing their content) in the reforming process, is shown in Figure 7 for the different experimental configurations. It can be observed that for percentages of oxygen lower than 2%, the CO2 tendency in the nonreformed and the reformed gas is opposite Figure 6. Time evolution of CO content in reformed gas (with Ni/γAl2O3catalyst) for different oxygen percentages (○0% O2,( ■)1% O2,( ●)2%O 2,▼6% O2and ▲10% O2) in the TZFBR regenerating gas stream. (Lines for visual help.) Figure 7. Methane (a), carbon dioxide (b), and C2hydrocarbons (c) content in nonreformed gas (■) and in reformed gas (○) for different reaction configurations, including several oxygen contents (0−10%) in regenerating gas fed to the TZFBR. (Lines for visual help.) Energy & Fuels Article dx.doi.org/10.1021/ef400227z |Energy Fuels XXXX, XXX, XXX−XXXG Reprinted with permission from Gas Catalytic Upgrading In A Two Zone Fluidized Bed Reactor Coupled To A CoGasification Plant, G. García, E. Campos, I. Fonts, J. L. Sanchez, and J. Herguido, Energy Fuels, DOI: 10.1021/ ef400227z , Publication Date (Web): 18 Apr 2013). Copyright (2013) American Chemical Society to that shown in Figure 5 for CO and H2. This trend may be expected as a result of reforming reactions. The decrease in CO2,CH 4and C2Hn, and the significant increase in the proportion of CO would evidence that the mass of C entering the TZFBR remains stable and therefore no coke is formed. The CO2and C2Hncontent decreases in the gas stream exiting from the dolomite bed outlet, and the presence of ethane + ethylene + acetylene is reduced to very low levels (lower than 0.018%, being around 1.1% after passing through the gasifier). On the other hand, the presence of CH4and CO increases significantly. These results are consistent with the recognized behavior of dolomite as a suitable catalyst for cracking and reforming heavy hydrocarbons. 49 As expected, according to the catalytic tests (Figure 4) the use of alumina in the TZFBR has no significant effect on the depletion of CO2and CH4, although in this case the presence of steam in the processed stream is likely promoting water−gas shift and other gasification and reforming reactions. However, C2hydrocarbons are almost completely removed with this configuration. As commented on previously, the differences between the percentages of CH4obtained in the experiment carried out only with dolomite and in the nonreformed stream of the experiments performed with the TZFBR could be caused by the deviations of the stoichiometric ratios, which varied between 28.8 and 32.3%. The use of Ni/γ-Al2O3catalyst reduces the presence of CH4, CO2, and C2Hnin the gas. On the one hand, the gas content in C2hydrocarbons is null, irrespective of the oxygen percentage used in the regenerating gas. On the other hand, its content in CH4and CO2reaches minimal values in the experiments carried out with 0 and 1% oxygen (e.g., CH4was not detected by GC during the experiment) but in experiments using 2, 6, and 10% oxygen in the gas, the CH4and CO2proportions increased. Thus, up to 2% O2, the CO2percentage in the gas is smaller than in the nonreformed gas, so no net gas oxidation occurs, and probably only coke is removed. Although this also leads to CO2formation, the increase in other compounds due to the catalytic reforming compensates for this. When 6 or 10% O2in the regenerating gas is used, the CO2in the reformed gas is higher than in the nonreformed stream due to the oxidation of the fuel species such as C2Hn,CH 4, and CO. The overall effect of all the tested experimental conditions on the quality of outgoing gas from the process can be analyzed in terms of gas LHV (see Figure 8). As CH4and CO increase while CO2decreases when dolomite is used, its heating value increases from 2590 kJ/m3(STP) (Table 5) to around 2800 kJ/ m3(STP), that is to say around 8%. In order to assess the effect of the catalytic treatment carried out in the TZFBR, the LHV of the gases upgraded in this reactor has been compared with the heating value of the nonreformed gas streams obtained in the same experiments since, as mentioned previously, the different stoichiometric ratios may cause a variation in the gas composition entering the TZFBR. Once the gas has undergone the dolomite treatment, alumina in the TZFBR causes a 28% increase in the LHV over the nonreformed stream, as can be seen in Figure 8. This increase is caused by the enrichment in CO and H2, as a consequence of the reforming reactions. The configuration with Ni/γ-Al2O3catalyst in the TZFBR also provokes a growth in the gas LHV when using 0 or 1% O2in the regenerating gas, but this entails working in a nonstationary state (Figure 6). Further increases in the oxygen supply above 2% cause the LHV to diminish, there being a sharp drop at 10% O2. The composition of the gas is downgraded since there is increasing gas combustion in the TZFBR, with a consequent decrease in CO and H2percentages in the gas products (Figure 5). An optimal condition is achieved working with 2% O2,as both stability and LHV upgrading over the nonreformed gas are achieved. Summing up, small percentages of O2(2% in this case) cause a slight increase in the CO2but also favor the cracking of heavier compounds, leading to a slight improvement in the gas heating value (see Figure 8). Using 2% of O2also ensures a stable operating process and aids the autothermicity of the process. Lastly, in order to apply this upgrading process in a full scale unit, it should be taken into account that pure (or high purity) oxygen should be used, mixed with a carrier agent that may be easily removed as steam that can be condensed prior to the final use of the produced gas. Moreover, the presence of a higher concentration of steam, which is a reforming agent, can be advantageous as the reforming reactions and hydrogen formation would be enhanced. 3.4. H2S Abatement using TZFBR. Regarding the H2S content in the produced gas (Figure 9), the use of dolomite allows its removal with an efficiency of 80 to 90%, depending on the experiment considered. There are differences in the H2S Figure 8. LHV percent increase percentage in TZFBR reformed gas regarding the nonreformed one, working with different solids (alumina and Ni/γ-Al2O3catalyst) and oxygen percentages (0−10%). Figure 9. H2S content in nonreformed gas and in reformed one in nonreformed gas (■) and in reformed one (○) for different reaction configurations, including several oxygen percentages (0−10%) in regenerating gas fed to the TZFBR. (Lines for visual help.) Energy & Fuels Article dx.doi.org/10.1021/ef400227z |Energy Fuels XXXX, XXX, XXX−XXXH Reprinted with permission from Gas Catalytic Upgrading In A Two Zone Fluidized Bed Reactor Coupled To A CoGasification Plant, G. García, E. Campos, I. Fonts, J. L. Sanchez, and J. Herguido, Energy Fuels, DOI: 10.1021/ ef400227z , Publication Date (Web): 18 Apr 2013). Copyright (2013) American Chemical Society percentage in the gas for experiments carried out with alumina in the TZFBR and for the rest carried out with Ni/γ-Al2O3 catalyst in this reactor. The H2S concentration for the nonreformed stream is between 200 and 300 ppm. The H2S leaving the dolomite bed is retained by the Ni/γ-Al2O3catalyst, resulting in its almost complete elimination. However, the total amount retained on the catalyst surface was not large enough to deactivate it in the runs carried out in an experimental series (e.g., those runs included in Figures 5−9), since all the runs were performed with the same sample of catalyst. In fact, taking into account the mass of catalyst placed in the TZFBR (42 g) and its Ni load (4.8 wt %), 1.16 g of H2S could be retained by the catalyst in the form of NiS. Moreover, if the gas flow fed to the TZFBR (0.3 dm3(STP)/min) contains at the most 300 ppm of H2S, the catalyst deactivation by sulfur poisoning will take place after around 8500 min of operation. In order to analyze more in depth the effects of H2Son catalyst deactivation and the subsequent outgoing gas composition, a final experimental configuration was adopted using Ni/γ-Al2O3catalyst in the TZFBR. In this new arrangement, the dolomite guard bed was omitted with the aim of exposing the catalyst to higher tar content in the gasification gas (avoiding the cracking effect of the dolomite). During a standard 90-min experiment, the system was stable and capable of cracking tars, attaining nondetectable tar content in the reformed gas as against 15 g tar/m3(STP) in the nonreformed gas. The variation in the composition of the TZFBR reformed gas compared with the nonreformed gas both including and not including dolomite bed is shown in Figure 10. As can be observed, in the configuration without the dolomite bed the H2S was still fully removed from the gas stream. This involves its removal for more than 90 min timeon-stream (i.e., more than 0.23 H2S mol/Ni mol in catalyst), consistent with the high selectivity of Ni catalyst found for hydrogen sulfide chemisorption. 27 Although a large amount of sulfur was adsorbed by the sites, the catalyst bed was still capable of maintaining its reforming activity of light hydrocarbons such as C2Hn, as occurred in the catalytic activity tests made with alumina only and described in section 3.1. However, methane was not reformed as a result of nickel poisoning with hydrogen sulfide (not retained by the dolomite this time). This behavior agrees with that observed previously corresponding to an active alumina support not deactivated by sulfur (Figure 4). As a consequence, a decrease in CO2, consumed in cracking reactions, and an increase in H2and CO, produced in reforming reactions, is perceived. Thus, the variation in the gas composition induced by the Ni/γ-Al2O3catalyst when no dolomite bed is used in the experimental arrangement is very similar to that resulting from the alumina activity test (Figure 4). 3.5. Characterization of the Used Catalyst. Figure 11 shows the results of the thermogravimetric oxidation analyses carried out on samples of the catalysts used in the runs using 0%, 1% and 6% of O2. As can be observed, the samples underwent a weight reduction in the temperature interval from 30 to 300 °C. The higher the oxygen percentage fed to the TZFBR in the previous experiment, the lower was the weight loss. In any case, this weight loss was not due to coke combustion, as evidenced by the fact that no CO2was detected by online mass spectrometry in the exhaust gas exiting from the TG equipment. Indeed, this can be observed in the MS signal (mass 44 Da) shown in Figure 11, corresponding to the catalyst from the 0% oxygen run in the TZFBR. In the range from 300 to 500 °C, a weight increase is perceived for all the solid samples. The higher the oxygen percentage fed to the TZFBR in the previous experiment, the lower was the weight gain. This behavior is related to nickel oxidation (Ni to NiO) and evidence the likely lower oxidation state of the solid when low oxygen percentages are used in the TZFBR. The original blue color of the solid samples corresponding to lower oxygen percentages (0% and 1%) turned to black at the end of the experiment. Moreover, in TG oxidation analysis, a peak of CO2 was detected by MS in the temperature range from 300 to 600 °C, evidencing the existence of coke depositions over the catalyst surface. For the 0% O2sample, this peak corresponds to a carbon content of approximately 0.03 mg C/gram of catalyst (see Figure 11). 4. CONCLUSIONS The main aim of this study is to upgrade the gasification gas obtained by the cogasification of sewage sludge, bituminous coal, and lignite. In the experimental study, a slip stream of the gas produced at lab scale was fed into a new catalytic reactor, the two-zone fluid bed reactor, which allows both the catalytic upgrading of the gas with Ni/γ-alumina catalyst and the in situ Figure 10. Composition variations in TZFBR reformed gas regarding the nonreformed one for both including and not including dolomite bed. Figure 11. Weight change and CO2mass spectra evolution of exhaust gases during the oxidation with air in temperature ramp of catalysts samples after experiments carried out with (●)%O 2,( ○)1%O 2, and (☆)6%O 2in the regenerating gas fed to the TZFBR. Energy & Fuels Article dx.doi.org/10.1021/ef400227z |Energy Fuels XXXX, XXX, XXX−XXXI Reprinted with permission from Gas Catalytic Upgrading In A Two Zone Fluidized Bed Reactor Coupled To A CoGasification Plant, G. García, E. Campos, I. Fonts, J. L. Sanchez, and J. Herguido, Energy Fuels, DOI: 10.1021/ ef400227z , Publication Date (Web): 18 Apr 2013). Copyright (2013) American Chemical Society regeneration of the catalyst deactivated by coke deposition. The use of this system enables a stable gasification gas stream to be achieved free of tars and H2S, one of the main contaminants in raw cogasification gas. The experiments included in this work started with a simple cogasification process, to which two other processes (desulfurization and gas reforming in TZFBR) were subsequently added. Each of these stages resulted in an added value to the technical features of the process under study. With the cogasification unit alone, a gas was obtained with relative low contents of H2and CO (7.9% and 9.1% respectively, volume basis), a LHV of 2.6 MJ/m3(STP) and high tar (15 g tar/m3(STP)) and H2S (1100 ppm) contents. When the desulfurization step (with dolomite) was added, the H2S content was reduced to less than 300 ppm, and the C2Hnand tars were to a large extent reformed and cracked. For example, the amount of naphthalene, the major compound in the produced tar and the only peak present in the chromatograms, in the gas generated after the gas passes through the dolomite guard bed, was smaller than 2 mg/m3(STP). Lastly, the complete system (cogasification + desulfurization + TZFBR) designed and started up in this work was capable of (1) removing almost the totality of hydrogen sulfide from the gas flow, (2) reducing tars in the gasification gas down to nondetection limits by GC-MS, with even the disappearance of the naphthalene peak, (3) reforming the gasification gas, improving its quality as a combustible gas, in terms of higher H2 and CO content, (4) increasing by 37% the LHV of the product gas, and (5) operating in a stable manner during continuous operation with a real gasification gas, while adjusting the relative oxygen flow necessary to regenerate the catalyst, the optimum oxygen percentage in the regeneration gas flow being around 2%. ■AUTHOR INFORMATION Corresponding Author *E-mail: [email protected]. 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Energy & Fuels Article dx.doi.org/10.1021/ef400227z |Energy Fuels XXXX, XXX, XXX−XXXK ef400227z , Publication D Reprinted with permission from Gas Catalytic Upgrading In A Two Zone Fluidized Bed Reactor Coupled To A CoGasification Plant, G. García, E. Campos, I. Fonts, J. L. Sanchez, and J. Herguido, Energy Fuels, DOI: 10.1021/ ef400227z , Publication Date (Web): 18 Apr 2013). Copyright (2013) American Chemical Society