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Catalytic cracking of raw bio-oil under FCC unit conditions over different zeolite-based catalysts
Journal of Industrial and Engineering Chemistry 78 (2019) 372–382
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Catalytic cracking of raw bio-oil under FCC unit conditions over different zeolite-based catalysts Álvaro Ibarra, Idoia Hita* , Miren J. Azkoiti, José M. Arandes, Javier Bilbao Chemical Engineering Department, University of the Basque Country (UPV/EHU), P. O. Box 644, 48080 Bilbao, Spain
A R T I C L E I N F O
A B S T R A C T
Article history: Received 30 January 2019 Received in revised form 12 May 2019 Accepted 20 May 2019 Available online 28 May 2019
The performance of different zeolite-based catalysts (HY, HZSM-5 and HBeta) on the catalytic cracking of bio-oil has been explored, using a simulated riser reactor and resembling industrial FCC conditions. The effect of the C/O (catalyst/bio-oil) ratio and the zeolite types have been assessed. The level of deoxygenation is >61% (increasing with C/O ratio). Total hydrocarbon yield was higher for the HBeta catalyst (56 wt%), while the liquid hydrocarbons yields were relatively similar for all catalysts, obtaining higher gasoline yields with the HY catalyst (46–55 wt%), and higher LPG yields with the HZSM-5 catalyst (12–14 wt%) due to its higher acidity. The HY zeolite produced more coke (4–7 wt%) given its capacity for retaining coke precursors within its micropores. © 2019 The Korean Society of Industrial and Engineering Chemistry. Published by Elsevier B.V. All rights reserved.
Keywords: Raw bio-oil FCC Catalytic cracking Zeolite Gasoline Deoxygenation
Introduction Technological development from the different routes of biomass valorization, to obtain fuels and raw materials, has contributed for reducing fossil fuel consumption and regulate the emission of greenhouse gases [1–4]. Fast pyrolysis from lignocellulosic biomass stands out from valorization routes, because it is a process with low environmental impact, possibility of scale-up, and their liquid product, bio-oil, has good prospects for large scale valorization [5–7]. Furthermore, fast pyrolysis can be operated in geographically delocalized units, and bio-oil can be transported (with a lower transport cost than biomass due to its higher energetic density) to refinery units carrying out biorefinery activities for the production of fuels with similar characteristics to fossil-originated ones [8,9]. Raw bio-oil is a very complex mixture of oxygenates (such as acids, alcohols, aldehydes, esters, ketones, phenols, guaiacol, syringol, sugars, furans) with a significant water content of 15– 35 wt% (dependent on the biomass drying process). In contrast with conventional fuel oils, bio-oil has a low heating value and volatility, a high thermal and chemical instability, and a corrosive nature [10–12]. For these reasons, prior to its valorization a previous bio-oil stabilization stage is recommended by adding a solvent (methanol) [13,14], thermal aging [15–17], or catalytic
* Corresponding author. E-mail address: [email protected] (I. Hita).
treatment (esterification, aldol condensation or mild hydrodeoxygenation), aiming for the removal of more unstable phenolic and acid compounds [18–20]. Aiming for the production of liquid biofuels from bio-oil in a great scale, two main well-established existing refinery routes have been predominantly explored: catalytic cracking [21,22], and hydroprocessing [20,23]. Besides the traditional vacuum gasoil (VGO), the fluid catalytic cracking (FCC) units are versatile for processing unconventional feeds, like secondary refinery streams [24,25], residual streams derived from the consumer society (i.e. plastic wastes) [26,27], or bio-derived streams like vegetable oil or bio-oil [21,28]. To date, catalytic cracking of bio-oil has attracted great research attention, although most of the available studies are focused on model compounds [29–32], or on conditioned bio-oil (i.e. through aging or mild hydrotreatment) [33–35], while studies on raw bio-oil are currently lagging behind. Additionally, many of these studies were conducted in different conditions to those of the FCC unit, generally at lower temperature and higher reactant contact time. Different authors have studied the cracking of model oxygenate compounds for bio-oil [22,32], mixtures of oxygenates with (VGO) [31,35–37], thermally aged bio-oil [38], bio-oil obtained from catalytic pyrolysis (partially cracked and deoxygenated) [39] and of the aqueous fraction of bio-oil (which does not contain the majority of the phenolic compounds) [40]. In cracking studies of bio-oil with VGO, it was determined that co-feeding up to 15–20 wt % bio-oil in the mixture, product yields were similar to those of the cracking of VGO alone, with an almost negligible yield of
https://doi.org/10.1016/j.jiec.2019.05.032 1226-086X/© 2019 The Korean Society of Industrial and Engineering Chemistry. Published by Elsevier B.V. All rights reserved.
A. Ibarra et al. / Journal of Industrial and Engineering Chemistry 78 (2019) 372–382
oxygenates [41]. However, when 35 wt% of bio-oil is co-fed, the results are different, with a decrease in gasoline yields (from 39 to 42 wt% in VGO cracking down to 36–38 wt% for the mixture). The different zeolite families (hydrated aluminosilicates), both natural and synthetic, play a key role in the sustainable industry given their interesting properties, like: high specific surface, shape selectivity features, acidity and basicity, hydrophility and hydrophobicity, hydrothermal stability, mechanical resistance and ecofriendly nature. Consequently, the synthesis of new zeolites and the modification of existing ones constitutes a broad subject for study in terms of their applications in heterogeneous catalysis, photocatalysis, gas separation and adsorption, water treatment, electrochemistry, pharmacy, etc. [42–46]. In the FCC unit, commercial catalysts are prepared based on a HY zeolite agglomerated in a silico-alumina and caolin matrix [47]. The HY zeolite presents a microporous structure suitable for the selective C5-C12 hydrocarbons, corresponding to the gasoline pool. Another key property of the catalyst is its hydrothermal stability, necessary in order to recover its activity after the subsequent reactionregeneration cycles for coke combustion [48]. Ibarra et al. [21,49] co-processing raw bio-oil (20 wt %) with VGO over commercial spent FCC catalyst (equilibrated in an industrial unit) have observed an interesting synergy between the cracking mechanisms of oxygenates in bio-oil and hydrocarbons in VGO, also stressing out the fact that an increase of the water concentration contributes to inhibiting undesired decarbonylation and decarboxylation reactions of oxygenates and attenuating coke formation and consequently catalyst deactivation. This manuscript delves into the catalytic cracking of raw bio-oil in similar conditions to those of the FCC unit, using three different catalysts: a commercial (HY zeolite-based) equilibrated FCC catalysts, and two agglomerated catalysts based in HZSM-5 (commonly used as additive for FCC catalysts) and HBeta (not typically used in FCC units and with intermediate properties compared to the other two zeolites). The goal of this study is to determine the effect of the different shape selectivity and acidity of the three zeolites in product distribution and composition, evaluating the quality of the product liquid as a fuel. The utilization of raw bio-oil as feed has also allowed for assessing the reactivity in FCC conditions of the different oxygenated families in raw bio-oil. Experimental Raw bio-oil The raw bio-oil feed was provided by Ikerlan/IK-4 (Vitoria, Spain), and was produced by fast pyrolysis of pine sawdust (450– 470 C) in a pilot plant (with a capacity of 25 kg h1) provided with a conical spouted bed reactor (CSBR) [50]. The raw bio-oil has been characterized by simulated distillation in an Agilent 6890 Series GC System equipped with a Simdis D2887 Fast/Ext. column, and elemental analysis (CHN, O content was determined by difference) using a LECO TruSpec CHN Macro apparatus. The water content in the bio-oil was quantified by Karl–Fisher titration in a Metrohm KF Titrino Plus 870. The concentrations of the different chemical families in the raw bio-oil were determined through gas chromatography combined with mass spectrometry (GC–MS) in a Shimadzu GC–MS QP2010 provided with a BPX5 column (50 m 0.22 mm 0.25 mm). Considering that bio-oil is known for its instability upon storage [51], the bio-oil used in this study was stored under a refrigerating environment at 4 C. GC–MS analyses were also performed at different stages of the experimental phase to check any possible compositional variations, but none were detected. The main properties of the raw bio-oil feed are summarized in Table 1, where the values provided are the average of three
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Table 1 Elemental analysis and chemical composition of the raw pine bio-oil. Elemental composition (wt %) C H S O Water content (wt %) Chemical composition (wt %) Acids and esters Acetic acid Aldehydes Hydroxy-acetaldehyde Ketones 1-Hydroxy-2-propanone Phenols Alcohols Ethers Sugars Levoglucosan Non identified
56.8 6.2 – 36.9 34.8 20.1 9.8 21.5 16.2 15.0 6.7 6.5 3.9 1.9 28.8 22.8 2.3
different GC MS analyses. The elemental analysis shows a composition of 56.8 wt% C, and an O content of 36.9 wt%, while on the other hand the water content was quantified in a 34.8 wt%. The main chemical groups present in the bio-oil were oxygenates, among of which sugars (28.8 wt%, mainly levoglucosan), aldehydes (21.5 wt%) and acids and esters (20.1 wt%) outstand. Catalysts Three zeolite-based catalysts with different acidity and shape selectivity features have been explored in this work: an equilibrated industrial FCC catalyst (equilibrated in reactionregeneration cycles), directly sampled from an FCC unit from Petronor S.A (Somorrostro, Spain) containing 15 wt% HY zeolite (designated as CY), and other two catalysts based on HZSM-5 and HBeta zeolites (supplied by Zeolyst International), designated as CZ and CB, respectively. The HBeta zeolite (CP814E, SiO2/Al2O3 = 25) and HZSM-5 zeolite (SiO2/Al2O3 = 80) have been activated to acquire their protonic form through calcination (5 C min1 ramp up to 550 C under a N2 stream for 4 h). After that, agglomeration of the HBeta and HZSM-5 zeolites in a matrix was carried out by wet extrusion of the active phase (zeolite, 20 wt%), binder (bentonite, 2 wt%) and inert filler (α-alumina, 78 wt%). Through catalyst agglomeration in a meso- and macroporous matrix, a hierarchical porous structure is formed, within of which sequential diffusion and partial cracking of bigger molecules from the feed takes place, followed by further cracking of the resulting chains inside the micropores, with a selectivity that is conditioned by zeolite acidity and shape selectivity. Two different types of perpendicular elliptic channels define the structure of the HZSM-5 zeolite: straight channels of 0.53 0.56 nm and sinusoidal ones of 0.51 0.55 nm, without cages in the intersections. On the other hand, HBeta zeolite consists of channels of 0.56 0.56 nm and 0.76 0.64 nm, forming openings of 0.56 0.65 nm, while the HY zeolite presents a structure with three-directional channels 0.74 nm wide which intersect forming cages (cavities) of a diameter of 1.24 nm [52,53]. In addition, the presence of the matrix in the catalyst particle facilitates the diffusion of coke precursors and attenuates the blockage of zeolite micropores. Subsequent to agglomeration, the catalysts were dried, calcined at 550 C, milled and sieved to the desired particle size (60120 mm). Finally, the prepared CZ and CB catalysts were equilibrated through a steaming treatment in a fluidized bed, for 5 h at 760 C and atmospheric pressure [54], aiming for simulating the equilibrating of the FCC catalyst. This
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treatment causes partial dealumination of the zeolites, removing strongly acidic centers from the catalytic surface, which are highly unstable and not capable of fully recovering their activity after successive reaction-regeneration stages [55]. Steaming treatments will also enhance the mesoporous structure of the catalysts, thus increasing the average pore diameter and enabling the accessibility of heavy bio-oil components into the zeolite micropores [56]. The physical properties, BET surface area and pore structure, were determined by N2 adsorption-desorption at 196 C in a Micromeritics ASAP 2010 unit, after a previous degassing of the sample at 150 C for 8 h in vacuum conditions. The crystalline structure was studied through X-ray diffraction in a Philips PW 1710 apparatus. Acidic properties (total acidity and strength of acidic sites) were determined by isothermal adsorption of NH3 at 150 C and a subsequent temperature programmed desorption (TPD) up to 550 C following a 5 C min1 ramp in a Setaram SDT 2960 thermobalance connected online with a Thermostar Balzers Instruments mass spectrometer. The Brönsted/Lewis (B/L) ratio of acidic sites was determined by FTIR spectrophotometry with pyridine adsorbed (Thermo Nicolet 740). Table 2 summarizes the main physico-chemical properties of the fresh catalysts. It was observed that due to the equilibration treatment applied to the catalysts (in the industrial unit that of the FCC catalyst), the difference in total acidity (in the 30–50 mmol gNH31 interval) and the average acidic strength (78–100 kJ mol1) are relatively small. However, there is a strong impact of these properties on the results, as will be discussed later on this manuscript. In addition, the crystallinity of the catalysts has been analyzed through high-angle X-ray diffraction (XRD) techniques in a Philips XPert PRO automatic diffractometer operating at 40 kV and 40 mA, in the u –2u configuration (Bragg Brentano geometry), with Cu Kα (l = 1.5418 Å). Data were collected from 3.5 to 60 (step size = 0.02, time per step = 100 s, total time =8 h). The observed spectra are characteristic and well-established in the literature for the HZSM-5 [57,58], HBeta [59] and HY [60,61] zeolites. Reaction equipment and product analysis The catalytic cracking runs have been carried out in a CREC (Chemical Reactor Engineering Centre) riser simulator designed to operate in laboratory scale under similar conditions than FCC unit reactor [62]. A detailed schematization of the FCC simulator is provided in Scheme 1 [21]. The operating conditions, resembling those typically used in industrial units, were: temperature, 550 C; contact time, 6 s; catalyst/bio-oil mass ratio (in a dry base), 1.8–7.3 gcatgbio-oil1. Please note that, in our experimental approach, the
Table 2 Properties of the fresh zeolite-based catalysts. Physical properties
CY
CZ
CB
SBET (m2 g1) Micropore surface (m2 g1) Pore volume (cm3 g1) Micropore volume (cm3 g1) dp (Å) Unit cell size (Å) Chemical properties Zeolite percentage (wt %) Total acidity (mmol g1) Average acid strength (kJ mol1) B/L ratio (mol mol1)
122 96 0.15 0.04 117.3 24.30
143 36 0.24 0.02 62.2 –
133 32 0.32 0.02 89.6 –
15 30 100 0.75
20 50 78 0.22
20 40 85 0.24
contact time simulates the time that, in an industrial FCC unit, the catalyst spends inside the riser section of the FCC unit, from the bottom entrance to the exit towards the regeneration stripping section. A fixed raw bio-oil mass of 0.11 g was upgraded on each experiment, and the employed catalyst mass varied between 0.13 and 0.53 g. Blank runs in no catalyst conditions (using only sand and an α-alumina + bentonite mixture, respectively) allowed for gathering insights into the thermal effect contribution and corroborate the inert nature of the matrix and binder used in catalyst preparation. The product stream was extracted from the reaction chamber once the reaction time was completed through a rapidly triggered vacuum pump connected to an auxiliary chamber (30 cm3), and then analyzed on-line by means of gas chromatography (GC) in an Agilent Technologies 7890 A chromatograph provided with flame ionization (FID) and pulsed flame photometric detectors (PFPD). In a pre-experimental phase, MS analyses were used to accurately identify the reaction products, while for routine product analysis GC techniques were only applied. The products were grouped in lumps, as follows: CO, CO2, dry gases (C1-C2), liquefied petroleum gases (LPG, C3-C4), gasoline (C5-C12), light cycle oil (LCO, C13-C20) and coke. The amount of coke was quantified from the mass loss after combustion of spent catalysts in a TGA-Q 5000 TA Instruments thermobalance, following a heating rate of 3 C min1 from 300 C up to 550 C. The rest of the volatile products was quantified through analysis of a gaseous sample in a Varian CP-4900 Micro GC, provided with two modules equipped with thermal conductivity detectors (TCD): (i) a molecular sieve (MS-5, 10 m) where H2, O2, N2, CH4 and CO are separated and (ii) a Porapak PPQ (10 m) where H2O, CO2 and light products (methane, ethylene, ethane, propylene, propane, acetaldehyde, butanes, butenes and acroleyn) are separated.
Scheme 1. Detailed schematization of the FCC unit riser simulator [21].
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The yield of each product fraction (Yi) was defined as specified by Eq. 1: Yi ¼
mi 100 mbiooil
ð1Þ
where mi, is the produced mass fraction of product i; and mbio-oil corresponds to the mass of bio-oil (wet basis) o the mass of oxygenates in bio-oil (dry basis). The conversion of each oxygenate component (Xi) of bio-oil was calculated from Eq. 2: Xi ¼
ðmi Þt¼0 ðmi Þt 100 ðmi Þt¼0
ð2Þ
where (mi)t=0 is the mass content of each oxygenate in the raw biooil feed and (mi)t the corresponding amount in the remaining biooil. Different parameters have been determined for evaluating the quality of the produced liquid fuel [11]. The deoxygenation degree (DOD, in wt%) was calculated from the oxygen mass content in the liquid fuel, (mO)fuel, and the oxygenates in bio-oil, (mO)feed, as: ðmO Þf uel DOD ¼ 1 ð3Þ 100 ðmO Þbiooil The liquid fuel yield (Yfuel, wt%) was defined as: Y f uel ¼
mf uel 100 mbiooil
ð4Þ
where mfuel is the mass of liquid fuel (hydrocarbons C5+ and oxygenates in the same boiling point range) and mbio-oil is the mass of bio-oil. On the other hand, the oxygen/carbon mass ratio (Ox/C, gOgC1) was determined from the oxygen mass content in the fuel product, (mO)fuel, divided by the carbon mass content in the fuel product, (mC)fuel, as: Ox ðmO Þf uel ¼ 100 C ðmC Þf uel
ð5Þ
Results and discussion
Fig. 1. Evolution with the catalyst/bio-oil mass ratio of the yields (on a wet basis) of (a) carbon products, (b) water and (c) coke, for the different catalysts.
Yields and product distribution The total yields of the main cracking products resulting from the dehydration of bio-oil (carbon volatile products and water) together with the yield of carbonaceous material deposited on the catalyst (coke) are displayed in Fig. 1. This Figure also displays the results of the thermal cracking (in a sand bed and without a catalyst) at C/O = 0, as well as the original carbon product and water content of the raw-bio oil. For the three studied catalysts, the yield of volatile carbon products (Fig. 1a) decreases upon increasing the catalyst/bio-oil ratio, while the yields of water and coke increase (Fig. 1b and c, respectively), as a direct consequence of higher catalyst/bio-oil ratios enhancing deoxygenation reactions (dehydration, decarbonylation and decarboxylation) when a higher amount of active catalyst sites become available. In addition, the C/ O ratio has a notable effect over coke yields because acid catalyst sites activate their formation mechanisms [63,64]. The results in Fig. 1 indicate that the reactions implied in bio-oil cracking are very rapid at 550 C, and thus become highly evolved even at the lowest catalyst/bio-oil ratio (1.8 gcatgbio-oil1), in such a way that, upon a further increase of the catalyst/bio-oil ratio, only slight variations are observed on dehydration and coke formation reactions. The yields of each product fraction in Fig. 1 are dependent on both the utilization of a catalyst and also the catalyst properties (mainly shape selectivity and secondarily catalyst acidity). This
way, the water yield trend (CZ > CB > CY, Fig. 1b) is in agreement with the total acidity trend (see Table 2) of the catalysts, known to enhance deoxygenation reactions [65]. Coke yields are overall very low in contrast to carbon volatile products (0–4 wt% vs 52–55 wt%, respectively). As observed in Fig. 1c, the CY catalyst is the one that yields the highest amount of coke (2.4–4.0 wt%). This can be attributable to its porous structure, containing cages formed in the intersections between the crystalline channels (which are not present in the structure of the HBeta and HZSM-5 zeolites), within of which coke precursors are trapped and further evolve towards polyaromatic structures, reaching similar coke yields to those obtained in the cracking of conventional FCC feeds [26]. It should be highlighted that, despite catalyst coking taking place, no operational drawbacks were observed upon feeding the bio-oil into the FCC simulator device. The fast injection in the setup, together with the significant water content in the bio-oil (known to prevent coke formation) and the severe temperature used in the process favor the fast vaporization of the feed and hence its fast conversion into desired products. On the other hand, the catalysts prepared from HZSM-5 and HBeta zeolites present more severe shape selectivity with narrower pores (see Table 2) which partially limit the formation of polyaromatics within the channels of the zeolite. Despite the well-known capacity of the HZSM-5 zeolite for minimizing coke
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formation, due to its microporous structure with channels that limit the progress of bimolecular hydrogen transfer and aromatic condensation reactions [53,66], the coke yields of the CZ catalyst (0.8–1.8 wt%) are overall higher than those of the CB catalyst (0.4– 1.0 wt%), which is attributable to the higher acidity of the CZ catalyst also playing a key role in coke condensation reactions [67]. The differences in lump product yields (in a dry basis) at different C/O ratios for the three catalysts are displayed in Fig. 2. The results at C/O = 0 correspond to a blank reaction (no catalyst and using sand). CO and CO2 formation reactions are favored upon increasing the catalyst/bio-oil ratio for all the catalysts (19.2– 27.6 wt%, Fig. 2a). CO and CO2 yields are higher for the CZ catalyst (22.5–27.6 wt%) as a consequence of decarboxylation and decarbonylation reactions being enhanced by catalyst acidity. On the other hand, the performance of the CY and CB catalysts resulted very similar (19.2–24.4 wt%) and, while the pore structure of the former (with wider average pores) enhances the diffusion for the oxygenates in bio-oil, the latter possesses a higher acidity that can outperform catalysts with overall wider pores. The compromise of the CB catalyst, in terms of oxygenate accessibility to the acidic sites (medium shape selectivity severity), and acidic sites themselves (intermediate acidity, Table 2), justifies its similar dry gas (14.8–19.5 wt%, Fig. 2b) and LPG (11.3–13.2 wt%, Fig. 2c) yields to those of the CZ catalyst (14.8–17.6 wt% and 12.1– 14.5 wt% for dry gas and LPG, respectively). On the other hand, these yields are lower for the CY catalyst (10.3–12.6 wt% dry gas
and 6.8–7.5 wt% LPG) due to its lower acidity and presumably more rapidly deactivated according to the results for coke yields (Fig. 1c, Fig. 2f). The gasoline yield (Fig. 2d), which resulted the main cracking product (34.0–55.3 wt%), follows the CY > CB > CZ trend, in opposite order to that of the total catalyst acidity. The lower acidity of the CY catalyst seems to be key for minimizing gasoline overcracking reactions towards dry gas and LPG, thus enhancing the yields of liquid fractions, which are also favored by operating at lower catalyst/bio-oil ratios. The velocity at which the transformation of bio-oil takes places is evidenced from the very low yields of LCO that were attained with all the catalysts (0.7–1.3 wt%, Fig. 2e), evidencing that the conversion of this fraction into lighter products is immediate, and negligible differences were observed between the three catalysts. This indicates that the different microporous structure of the corresponding zeolites does not have a strong impact in terms of diffusional limitations of LCO components. As discussed earlier on the results in Fig. 1, the higher coke yield (4.2–7.0 wt%, Fig. 2f) is attained with the CY catalyst due to the different porous structure of zeolite HY in comparison with the other two catalysts, lacking macropores in their matrix and without cages in their pore intersections, hence favoring the diffusion of coke precursors towards the outside of the catalyst. The lower coke yield is for the CB catalyst (0.7–1.7 wt%) thanks to a compromise between moderate shape selectivity and acidity in comparison with the other catalysts.
Fig. 2. Evolution with the catalyst/bio-oil mass ratio of the product yields (in a dry basis) for the different catalysts.
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In the catalytic cracking of bio-oil, considerable amounts of hydrocarbons are produced, from complex combinations of cracking, deoxygenation, dehydration, decarboxylation, decarbonylation, and alkylation reactions, among others. Fig. 3 shows the total hydrocarbon yields (in a dry basis) (Fig. 3a), and also those corresponding to the liquid C5-C20 fractions (Fig. 3b) that can be potentially blended in the gasoline and diesel pools. The higher total hydrocarbon yield corresponds to the CB catalyst due to the aforementioned compromise between its shape selectivity and acidity, being of 42.6 wt% at a catalyst/bio-oil ratio of 7.3 gcatgbio1 oil , and slightly higher than that corresponding to the CZ catalyst (40.2 wt%) at the same conditions. Regarding the C5-C20 fraction (Fig. 3b), the higher yields attained with the CB catalyst are clear, reaching a maximum value of ca. 14 wt% for a catalyst/bio-oil ratio in the 4–5 gcatgbio-oil1 range. At higher C/O values, overcracking reactions acquire greater relevance and, consequently, liquid hydrocarbon fractions are lower. Table 3 allows for a detailed evaluation of the effect of the catalyst/bio-oil ratio and the different catalysts on the yield of the dry gas (C1-C2) and LPG (C3-C4) fractions. The C1-C2 (dry gas) yields are overall lower for the CY catalyst, along the same lines of the yields reported in Fig. 2b and d for this catalyst. The comparison between the results obtained with the CZ and CB catalysts is intricate and, even though a lower methane yield (7.7–8.5 wt%) was attained with the CZ catalyst, the rest of the yields do not present significant differences in the studied C/O range. However, it was observed that for a catalyst/bio-oil ratio of 7.3 gcatgbio-oil1, the CB catalyst provides a higher ethylene yield (6.7 wt%), which might be of industrial interest, at the expense of a higher methane production. Regarding the LPG fraction (C3-C4 range), operating with the CY catalyst leads mainly to olefins like propylene (3.0– 4.3 wt%) and butenes (1.9–2.2 wt%). Notably, the yields of butanes are higher with the CB (1.1–2.4 wt%) and, particularly, with the CZ catalyst (2.7–4.2 wt%). An overall decreasing tendency was observed in terms of the olefinic nature of the individual C2, C3
Fig. 3. Evolution with the catalyst/bio-oil mass ratio of (a) the total hydrocarbon yield (in a dry basis) and (b) C5-C20 hydrocarbon yields for the different catalysts.
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Table 3 Gas product yields (wt%) at the lowest and highest studied catalyst/bio-oil ratios in the cracking of raw bio-oil over different zeolites.
Methane, CH4 Ethylene, C2H4 Ethane, C2H6 Propylene, C3H6 Propane, C3H8 Butenes, C4H8 Butanes, C4H10
Catalyst/bio-oil (gcatgbio-oil1)
CY
CZ
CB
1.8 7.3 1.8 7.3 1.8 7.3 1.8 7.3 1.8 7.3 1.8 7.3 1.8 7.3
5.7 6.9 3.0 3.6 1.6 2.0 4.3 3.1 0.2 1.0 2.2 1.9 0.1 1.6
7.7 8.5 4.9 5.9 2.0 3.0 5.6 4.8 0.8 0.9 3.1 2.7 2.7 4.2
8.2 9.7 4.7 6.7 2.1 3.1 5.1 5.8 0.9 1.3 4.2 3.7 1.1 2.4
and C4 fractions upon increasing the catalyst/bio-oil ratio as a consequence of hydrogen transfer reactions being favored due to the higher amount of available acidic catalyst sites. No clear differences are observed among the catalysts for the transformation of ethylene into ethane. On the other hand, in the propylene to propane transformation, and even more in that of butenes into butanes, the different catalyst acidity seems to acquire importance, explaining the CY > CB > CZ olefinity trend of the C4 fraction, with an opposite order to the acidity trend. Conversion of individual bio-oil oxygenates The effect of the catalyst/bio-oil ratio on the conversion of different individual oxygenate compounds (representative of the different chemical families in bio-oil) for the three catalysts is displayed in Fig. 4. These compounds are: acetaldehyde (from lineal aldehydes), furfural (from furans), acetic acid (from acids), 3methyl-1,2,cyclopentenone (from cyclic ketones), hydroxyacetone (from lineal ketones), phenol (from branched phenolics), levoglucosan (from high molecular weight sugars), and phenolic compounds (to also consider the results corresponding to phenolic ethers, like guaiacol, syringol, and their derivates). The catalytic cracking mechanisms for biomass-derived oxygenates is wellestablished and, as reported by Huber and Corma [1], hydrogen producing and consuming reactions are of critical importance. The conversions of acetaldehyde (Fig. 4a), hydroxyacetone (Fig. 4e) and levoglucosan (Fig. 4h) are considerably high even for a catalyst/bio-oil ratio of 1.8 gcatgbio-oil1 for all the catalysts. The high acetaldehyde conversion might imply its partial undesired transformation into coke, given the important role that this component plays in the reactions involved in the formation of pyrolytic lignin due to the formation of trimethyltrioxane as intermediate [29]. On the other hand, the high conversion of levoglucosan (present in high proportion in bio-oil derived from fast pyrolysis) can be explained from its high instability and reactivity [68]. On the contrary, the conversion of furfural (Fig. 4b) and 3-methyl-1,2-cyclopentenone (Fig. 4d) is lower at the studied conditions, increasing up to almost full conversion (94.7 wt%) with the most active CZ catalyst for a catalyst/bio-oil of 7.3 gcatgbio-oil1. The conversion for acetic acid (Fig. 4c), phenol (Fig. 4f) and the phenolic groups in bio-oil (Fig. 4g) is lower and, even if it increases upon increasing the catalyst/bio-oil ratio, it is still far from reaching full conversion, with acetic acid being the least reactive component. Horne and Williams [69], reporting studies of furfural cracking over HZSM-5 zeolites, mainly obtained aromatic products (mostly alkylbenzenes) due to the higher acidity of their catalyst and the high space time, which facilitates olefin condensation. Bertero et al. [32]. using a fixed bed MAT reactor at
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Fig. 4. Evolution with the catalyst/bio-oil ratio of individual conversions of the different oxygenated families in the cracking of raw bio-oil over different catalysts.
500 C and 60 s of time on stream, and an equilibrated FCC catalyst, produced majoritarily pentenes from the cracking of furfural and aromatics from the cracking of 3-methyl-1,2-cyclopentenone. The reactivity order of the bio-oil oxygenates in the reaction media (corresponding to raw bio-oil), is in agreements with the results reported in the literature for the cracking of these components individually feed (at different conditions and using different catalysts) [30,29,22,70,71]. When comparing the results for the different catalysts, the conversion order, with the exception of phenol and phenolic compounds, is the following: CZ > CB > CY, which corresponds to the acidity trend of the catalysts. These results indicate that the total acidity of the used catalyst will have a greater impact on the conversion of individual oxygenates than their pore network structure.
This situation also takes place in the conversion of levoglucosan, a voluminous unstable molecule, whose conversion is significant at the studied temperature even without a catalyst [71], and almost complete at low catalyst/bio-oil ratios (Fig. 4h). However, the conversion of levoglucosan over the CZ catalyst is only slightly higher than over the CB catalyst, which seems to indicate that the HZSM-5 zeolite presents diffusion limitations (narrower average pores, see Table 2) for the levoglucosan molecule, which is favored by HBeta zeolite. In the conversion of phenol and phenolic groups, the conversion order does not follow the increasing acidity trend, and the CB catalyst appears as the most active, which points out a compromise between catalyst acidity and pore accessibility to acidic sites in the case of this catalyst. The low activity of the CY catalyst should be highlighted in the conversion of these compounds, attributable to
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its lower acidity in fresh conditions as well as the fact that it might already be considerable deactivated at such short reaction time as 6 s, based on a previous work by Ibarra et al. [21].
Table 4 Quality indexes of the liquid cracking product of raw bio-oil obtained with the different catalysts at the lowest and highest studied catalyst/bio-oil ratios. Parameters
catalyst/bio-oil (gcatgbio-oil1)
CY
CZ
CB
DOD
1.8 7.3 1.8 7.3 1.8 7.3
61.1 70.7 33.1 27.2 0.50 0.38
70.7 80.4 26.7 19.7 0.38 0.31
67.7 77.3 29.8 22.1 0.40 0.34
Quality of the fuel Table 4 summarizes the indexes for assessing the quality of the liquid product from raw bio-oil cracking: deoxygenation degree (DOD, Eq. 3), liquid fuel yield (Yfuel, Eq. 4), and oxygen/carbon ratio (Ox/C, Eq. 5). The extent of the DOD follows the CZ > CB > CY trend, which highlights the fact that catalyst acidity features will not only heavily determine product distribution, but is also a key property for increasing oxygen removal by enhancing cracking and deoxygenation (dehydration, decarboxylation and decarbonylation) reactions of the most part of bio-oil components. All in all, the deoxygenation efficiency is high for the three catalysts (61.1–80.4 %) increasing with the catalyst/bio-oil ratio (higher availability of acidic sites for reactants). Specifically, using the CZ catalyst, DOD indexes from 70.7% to 80.4% were attained. However, a higher oxygen removal in the products will come at the expense of lower liquid fuel yields (Yfuel, 22.1–33.1 wt% in the studied C/O range) and higher coke and gas product yields (see results in Fig. 2) and, therefore, a compromise between the different operation parameters (process severity) and desired production goals must be established. An increase in the catalyst/bio-oil ratio also shows a tendency for slightly decreasing the Ox/C ratio (defined in Eq. 5) in the liquid product, which is a key quality index to evaluate the fuel. Upon decreasing the Ox/C ratio, the H/C ratio in the fuel is higher (from
379
Yliquid Ox/C
fuel
0.2 to 0.4 in the catalyst/bio-oil range of 1.8–7.3 gcatgbio-oil1 for the CY catalyst). It was observed that the lowest Ox/C value was attained at a catalyst/bio-oil ratio of 7.3 gcatgbio-oil1 and using the CY catalyst which, with its higher acidity, provides a high-quality liquid fuel with a reduced oxygen content (80.8% removal). However, as mentioned earlier, this also implies a lower liquid product yield (27.2 wt%) in contrast to lower catalyst/bio-oil ratios, due to excessive overcracking of the gasoline fraction towards dry gas and LPG. Yields of the gasoline fraction components The gasoline fraction consists of C5-C12 hydrocarbons and oxygenates in the same boiling point range and, therefore, assessing both fractions separately is advisable. Fig. 5 displays the yields corresponding to the main oxygenate and hydrocarbon families in the gasoline fraction for the different catalysts, where important effects of both the catalyst and catalyst/bio-oil ratio
Fig. 5. Comparison of the yields of hydrocarbons (graph a and b) and oxygenates (graph c and d) in the gasoline fraction obtained in the cracking of raw bio-oil over different catalysts at the lowest and highest studied catalyst/bio-oil ratios.
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were observed. For a catalyst/bio-oil ratio of 1.8 gcatgbio-oil1, overall, the main hydrocarbons are olefins (3.1–4.3 wt%), with lower yields of i-paraffins (2.1–3.3 wt%), aromatics (1.6–3.3 wt%), naphthenes (1.3–2.8 wt%) and n-paraffins (0.3–0.7 wt%), as seen in Fig. 5a, attaining total hydrocarbon yields of 10.1–12.2 wt%. An increase in the catalyst/bio-oil ratio up to 7.3 gcatgbio-oil1 (Fig. 5b) notably modifies the hydrocarbon composition in gasoline, with significantly lower olefin and naphthene yields. The former are mainly affected by enhanced condensation of aromatics (which increase up to 4.8 wt%), while the latter mostly undergo cracking reactions towards paraffins and olefins, which later on will likely also condense into aromatics. Interestingly, total hydrocarbon yields (11.9–12.3 wt%) are very similar to those obtained at lower catalyst/bio-oil ratio. At this high catalyst/bio-oil ratio the mildly acidic CB catalyst presents more advantages, with i-paraffins being the main chemical group (5 wt%), while the aromatic yield only increases slightly in comparison with the lowest catalyst/bio-oil value, accounting for only 50% of the yields attained with the other catalysts (2.3 wt% vs 4.7–4.8 wt%). Furthermore, the yield of olefins, even if higher than for the other catalysts, is still notably lower than at lower catalyst/bio-oil value. The higher acidity of the CZ catalyst in contrast to the CY catalyst had an enhancing effect over cracking and deoxygenation reactions, providing higher yields of hydrocarbons (with the exception of aromatic compounds). However, at the same time, the transformation of phenol and phenolic compounds is hindered within the pore structure of the HZSM-5 zeolite due to steric difussional limitations. Likewise, these limitations restrict Diels-Alder condensation reactions of olefins towards
aromatic compounds and, according to our results, also cracking reactions of naphthenes. It should also be considered that, since olefins are primary products, their formation is favored by catalyst acidity, which at the same time also enhances their transformation into n-paraffins and, in a lesser extent, into iparaffins. On the other hand, the better results achieved with the CB catalyst, among which lower aromatic and higher naphthene yields outstand, are attributable to a good compromise between moderate acidity and shape selectivity, which not only hinders olefin condensation into aromatics, but also olefin cracking (into dry gas and LPG) and naphthene cracking. On the other hand, the yields of oxygenates in the gasoline fraction are complementary results to those of the conversion of individual components in bio-oil, discussed in Fig. 4. As seen from Fig. 5c and Fig. 5d, oxygenated compounds constitute the main fraction in gasoline with yields in the 34.6–45.0 wt% and 21.9– 33.2 wt% range, at 1.8 and 7.3 gcatgbio-oil1, respectively. Upon increasing the catalyst/bio-oil ratio, gasoline becomes less oxygenated with deoxygenation being favored, as previously discussed from the results in Table 4, due to a higher availability of active sites. Additionally, it was observed that the least acidic CY catalyst provides higher oxygenate yields (with the exception of ketones), due to the lower activity of the CY catalyst in cracking and deoxygenation reactions. The yields of oxygenates decrease with more acidic catalysts. The lower yield of acetaldehyde and phenol (with the CZ and CB catalysts respectively) should be highlighted, since these results are relevant for minimizing issues derived from the condensation of phenolic components in bio-oil storage and handling.
Fig. 6. Comparison for different catalysts of the distribution of the yields of (a) aromatics, (b) n-paraffins, (c) olefins and (d) i-paraffins with different C atom number in the gasoline fraction. (Catalyst/bio-oil ratio = 1.8 gcatgbio-oil1).
A. Ibarra et al. / Journal of Industrial and Engineering Chemistry 78 (2019) 372–382
Delving into the results in Fig. 5, interesting differences were observed upon evaluating individual yields of each hydrocarbon family (on a carbon number basis) in the gasoline fraction, as seen in Fig. 6. At a low catalyst/bio-oil ratio (1.8 gcatgbio-oil1), the aromatic components detected are mainly in the C6-C8 range (Fig. 6a), and with notably lower yields for the CB catalyst. With respect to the comparison between the CY and CZ catalysts, steric hindrances are again evidenced, with the microporous HZSM-5 structure limiting the formation of aromatics, either through cracking of phenolics or olefin condensation. The yield of n-paraffins (Fig. 6b) is very low (<0.2 wt%), and therefore the measurements are liable to a higher experimental error. A heterogeneous distribution of the n-paraffins is observed in terms of carbon atoms, with both light and heavy paraffins being present, and which is also affected by catalyst properties. Olefins (Fig. 6c) are mainly C5-C6, with higher yields for the CB catalyst which, as previously discussed, has a limited capacity for transforming these olefins (either through aromatic condensation of light olefin cracking reactions). In the same way, C5 and C6 iparaffins are predominant (Fig. 6d), with higher yields for more acidic catalysts (CB, CZ), which indicates that the higher acidity is enhancing their transformation from olefins through hydrogen transfer reactions. Conclusions The shape selectivity and acidity of the zeolite in the catalyst have a significant impact on catalyst activity, selectivity and coke deposition on the cracking of raw bio-oil in FCC conditions. The commercial FCC (HY zeolite-based) catalyst is adequate for maximizing liquid fuel yields, with a higher gasoline yield and a limited gas product formation. At the same time, coke yield is favored in this catalyst due to the higher retention capacity of coke precursors. On the other hand, the more acidic HZSM-5-based catalyst is suitable for producing light olefins and the HBeta-based one is intermediate between the other two. The concentration of hydrocarbons in the gasoline fraction, with olefins and aromatics as the main compounds, is dependent on the porous structure and acidity of the zeolites. Consequently, the aromatic content is higher in the gasoline produced using the HY zeolite and lower with the HZSM-5 zeolite. Furthermore, the cracking results are dependant on the composition of bio-oil, due to the distinct reactivity of its components, in the following order: sugars (levoglucosan) >lineal aldehydes (acetaldehyde) >lineal ketones (hydroxyacetone) >furans (furfural) cyclic ketones (3-methyl-1,2-cyclopentenone) >acids (acetic acid) branched phenols (phenol) phenolic groups. Acknowledgements This work has been carried out with financial support from the Ministry of Economy and Competitiveness of the Spanish Government (MINECO) (Projects CTQ2012-35192, co-funded with FEDER funds) and the Basque Government (Project IT748-13). Dr. Idoia Hita is grateful for her postdoctoral grant awarded by the Department of Education, University and Research of the Basque Government (grant number POS_2015_1_0035). Dr. Alvaro Ibarra is grateful for his PhD grant from the University of the Basque Country (UPV/EHU). Authors also acknowledge Ikerlan/IK-4 and Petronor S.A. for providing the raw bio-oil and the catalyst used in this work, respectively. References [1] G.W. Huber, A. Corma, Angew. Chemie Int. Ed. 46 (38) (2007) 7184. [2] S. Gunukula, S.J.W. Klein, H.P. Pendse, W.J. DeSisto, M.C. Wheeler, Appl. Energy 214 (2018) 16.
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Catalytic cracking of raw bio-oil under FCC unit conditions over different zeolite-based catalysts Álvaro Ibarra, Idoia Hita* , Miren J. Azkoiti, José M. Arandes, Javier Bilbao Chemical Engineering Department, University of the Basque Country (UPV/EHU), P. O. Box 644, 48080 Bilbao, Spain
A R T I C L E I N F O
A B S T R A C T
Article history: Received 30 January 2019 Received in revised form 12 May 2019 Accepted 20 May 2019 Available online 28 May 2019
The performance of different zeolite-based catalysts (HY, HZSM-5 and HBeta) on the catalytic cracking of bio-oil has been explored, using a simulated riser reactor and resembling industrial FCC conditions. The effect of the C/O (catalyst/bio-oil) ratio and the zeolite types have been assessed. The level of deoxygenation is >61% (increasing with C/O ratio). Total hydrocarbon yield was higher for the HBeta catalyst (56 wt%), while the liquid hydrocarbons yields were relatively similar for all catalysts, obtaining higher gasoline yields with the HY catalyst (46–55 wt%), and higher LPG yields with the HZSM-5 catalyst (12–14 wt%) due to its higher acidity. The HY zeolite produced more coke (4–7 wt%) given its capacity for retaining coke precursors within its micropores. © 2019 The Korean Society of Industrial and Engineering Chemistry. Published by Elsevier B.V. All rights reserved.
Keywords: Raw bio-oil FCC Catalytic cracking Zeolite Gasoline Deoxygenation
Introduction Technological development from the different routes of biomass valorization, to obtain fuels and raw materials, has contributed for reducing fossil fuel consumption and regulate the emission of greenhouse gases [1–4]. Fast pyrolysis from lignocellulosic biomass stands out from valorization routes, because it is a process with low environmental impact, possibility of scale-up, and their liquid product, bio-oil, has good prospects for large scale valorization [5–7]. Furthermore, fast pyrolysis can be operated in geographically delocalized units, and bio-oil can be transported (with a lower transport cost than biomass due to its higher energetic density) to refinery units carrying out biorefinery activities for the production of fuels with similar characteristics to fossil-originated ones [8,9]. Raw bio-oil is a very complex mixture of oxygenates (such as acids, alcohols, aldehydes, esters, ketones, phenols, guaiacol, syringol, sugars, furans) with a significant water content of 15– 35 wt% (dependent on the biomass drying process). In contrast with conventional fuel oils, bio-oil has a low heating value and volatility, a high thermal and chemical instability, and a corrosive nature [10–12]. For these reasons, prior to its valorization a previous bio-oil stabilization stage is recommended by adding a solvent (methanol) [13,14], thermal aging [15–17], or catalytic
* Corresponding author. E-mail address: [email protected] (I. Hita).
treatment (esterification, aldol condensation or mild hydrodeoxygenation), aiming for the removal of more unstable phenolic and acid compounds [18–20]. Aiming for the production of liquid biofuels from bio-oil in a great scale, two main well-established existing refinery routes have been predominantly explored: catalytic cracking [21,22], and hydroprocessing [20,23]. Besides the traditional vacuum gasoil (VGO), the fluid catalytic cracking (FCC) units are versatile for processing unconventional feeds, like secondary refinery streams [24,25], residual streams derived from the consumer society (i.e. plastic wastes) [26,27], or bio-derived streams like vegetable oil or bio-oil [21,28]. To date, catalytic cracking of bio-oil has attracted great research attention, although most of the available studies are focused on model compounds [29–32], or on conditioned bio-oil (i.e. through aging or mild hydrotreatment) [33–35], while studies on raw bio-oil are currently lagging behind. Additionally, many of these studies were conducted in different conditions to those of the FCC unit, generally at lower temperature and higher reactant contact time. Different authors have studied the cracking of model oxygenate compounds for bio-oil [22,32], mixtures of oxygenates with (VGO) [31,35–37], thermally aged bio-oil [38], bio-oil obtained from catalytic pyrolysis (partially cracked and deoxygenated) [39] and of the aqueous fraction of bio-oil (which does not contain the majority of the phenolic compounds) [40]. In cracking studies of bio-oil with VGO, it was determined that co-feeding up to 15–20 wt % bio-oil in the mixture, product yields were similar to those of the cracking of VGO alone, with an almost negligible yield of
https://doi.org/10.1016/j.jiec.2019.05.032 1226-086X/© 2019 The Korean Society of Industrial and Engineering Chemistry. Published by Elsevier B.V. All rights reserved.
A. Ibarra et al. / Journal of Industrial and Engineering Chemistry 78 (2019) 372–382
oxygenates [41]. However, when 35 wt% of bio-oil is co-fed, the results are different, with a decrease in gasoline yields (from 39 to 42 wt% in VGO cracking down to 36–38 wt% for the mixture). The different zeolite families (hydrated aluminosilicates), both natural and synthetic, play a key role in the sustainable industry given their interesting properties, like: high specific surface, shape selectivity features, acidity and basicity, hydrophility and hydrophobicity, hydrothermal stability, mechanical resistance and ecofriendly nature. Consequently, the synthesis of new zeolites and the modification of existing ones constitutes a broad subject for study in terms of their applications in heterogeneous catalysis, photocatalysis, gas separation and adsorption, water treatment, electrochemistry, pharmacy, etc. [42–46]. In the FCC unit, commercial catalysts are prepared based on a HY zeolite agglomerated in a silico-alumina and caolin matrix [47]. The HY zeolite presents a microporous structure suitable for the selective C5-C12 hydrocarbons, corresponding to the gasoline pool. Another key property of the catalyst is its hydrothermal stability, necessary in order to recover its activity after the subsequent reactionregeneration cycles for coke combustion [48]. Ibarra et al. [21,49] co-processing raw bio-oil (20 wt %) with VGO over commercial spent FCC catalyst (equilibrated in an industrial unit) have observed an interesting synergy between the cracking mechanisms of oxygenates in bio-oil and hydrocarbons in VGO, also stressing out the fact that an increase of the water concentration contributes to inhibiting undesired decarbonylation and decarboxylation reactions of oxygenates and attenuating coke formation and consequently catalyst deactivation. This manuscript delves into the catalytic cracking of raw bio-oil in similar conditions to those of the FCC unit, using three different catalysts: a commercial (HY zeolite-based) equilibrated FCC catalysts, and two agglomerated catalysts based in HZSM-5 (commonly used as additive for FCC catalysts) and HBeta (not typically used in FCC units and with intermediate properties compared to the other two zeolites). The goal of this study is to determine the effect of the different shape selectivity and acidity of the three zeolites in product distribution and composition, evaluating the quality of the product liquid as a fuel. The utilization of raw bio-oil as feed has also allowed for assessing the reactivity in FCC conditions of the different oxygenated families in raw bio-oil. Experimental Raw bio-oil The raw bio-oil feed was provided by Ikerlan/IK-4 (Vitoria, Spain), and was produced by fast pyrolysis of pine sawdust (450– 470 C) in a pilot plant (with a capacity of 25 kg h1) provided with a conical spouted bed reactor (CSBR) [50]. The raw bio-oil has been characterized by simulated distillation in an Agilent 6890 Series GC System equipped with a Simdis D2887 Fast/Ext. column, and elemental analysis (CHN, O content was determined by difference) using a LECO TruSpec CHN Macro apparatus. The water content in the bio-oil was quantified by Karl–Fisher titration in a Metrohm KF Titrino Plus 870. The concentrations of the different chemical families in the raw bio-oil were determined through gas chromatography combined with mass spectrometry (GC–MS) in a Shimadzu GC–MS QP2010 provided with a BPX5 column (50 m 0.22 mm 0.25 mm). Considering that bio-oil is known for its instability upon storage [51], the bio-oil used in this study was stored under a refrigerating environment at 4 C. GC–MS analyses were also performed at different stages of the experimental phase to check any possible compositional variations, but none were detected. The main properties of the raw bio-oil feed are summarized in Table 1, where the values provided are the average of three
373
Table 1 Elemental analysis and chemical composition of the raw pine bio-oil. Elemental composition (wt %) C H S O Water content (wt %) Chemical composition (wt %) Acids and esters Acetic acid Aldehydes Hydroxy-acetaldehyde Ketones 1-Hydroxy-2-propanone Phenols Alcohols Ethers Sugars Levoglucosan Non identified
56.8 6.2 – 36.9 34.8 20.1 9.8 21.5 16.2 15.0 6.7 6.5 3.9 1.9 28.8 22.8 2.3
different GC MS analyses. The elemental analysis shows a composition of 56.8 wt% C, and an O content of 36.9 wt%, while on the other hand the water content was quantified in a 34.8 wt%. The main chemical groups present in the bio-oil were oxygenates, among of which sugars (28.8 wt%, mainly levoglucosan), aldehydes (21.5 wt%) and acids and esters (20.1 wt%) outstand. Catalysts Three zeolite-based catalysts with different acidity and shape selectivity features have been explored in this work: an equilibrated industrial FCC catalyst (equilibrated in reactionregeneration cycles), directly sampled from an FCC unit from Petronor S.A (Somorrostro, Spain) containing 15 wt% HY zeolite (designated as CY), and other two catalysts based on HZSM-5 and HBeta zeolites (supplied by Zeolyst International), designated as CZ and CB, respectively. The HBeta zeolite (CP814E, SiO2/Al2O3 = 25) and HZSM-5 zeolite (SiO2/Al2O3 = 80) have been activated to acquire their protonic form through calcination (5 C min1 ramp up to 550 C under a N2 stream for 4 h). After that, agglomeration of the HBeta and HZSM-5 zeolites in a matrix was carried out by wet extrusion of the active phase (zeolite, 20 wt%), binder (bentonite, 2 wt%) and inert filler (α-alumina, 78 wt%). Through catalyst agglomeration in a meso- and macroporous matrix, a hierarchical porous structure is formed, within of which sequential diffusion and partial cracking of bigger molecules from the feed takes place, followed by further cracking of the resulting chains inside the micropores, with a selectivity that is conditioned by zeolite acidity and shape selectivity. Two different types of perpendicular elliptic channels define the structure of the HZSM-5 zeolite: straight channels of 0.53 0.56 nm and sinusoidal ones of 0.51 0.55 nm, without cages in the intersections. On the other hand, HBeta zeolite consists of channels of 0.56 0.56 nm and 0.76 0.64 nm, forming openings of 0.56 0.65 nm, while the HY zeolite presents a structure with three-directional channels 0.74 nm wide which intersect forming cages (cavities) of a diameter of 1.24 nm [52,53]. In addition, the presence of the matrix in the catalyst particle facilitates the diffusion of coke precursors and attenuates the blockage of zeolite micropores. Subsequent to agglomeration, the catalysts were dried, calcined at 550 C, milled and sieved to the desired particle size (60120 mm). Finally, the prepared CZ and CB catalysts were equilibrated through a steaming treatment in a fluidized bed, for 5 h at 760 C and atmospheric pressure [54], aiming for simulating the equilibrating of the FCC catalyst. This
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treatment causes partial dealumination of the zeolites, removing strongly acidic centers from the catalytic surface, which are highly unstable and not capable of fully recovering their activity after successive reaction-regeneration stages [55]. Steaming treatments will also enhance the mesoporous structure of the catalysts, thus increasing the average pore diameter and enabling the accessibility of heavy bio-oil components into the zeolite micropores [56]. The physical properties, BET surface area and pore structure, were determined by N2 adsorption-desorption at 196 C in a Micromeritics ASAP 2010 unit, after a previous degassing of the sample at 150 C for 8 h in vacuum conditions. The crystalline structure was studied through X-ray diffraction in a Philips PW 1710 apparatus. Acidic properties (total acidity and strength of acidic sites) were determined by isothermal adsorption of NH3 at 150 C and a subsequent temperature programmed desorption (TPD) up to 550 C following a 5 C min1 ramp in a Setaram SDT 2960 thermobalance connected online with a Thermostar Balzers Instruments mass spectrometer. The Brönsted/Lewis (B/L) ratio of acidic sites was determined by FTIR spectrophotometry with pyridine adsorbed (Thermo Nicolet 740). Table 2 summarizes the main physico-chemical properties of the fresh catalysts. It was observed that due to the equilibration treatment applied to the catalysts (in the industrial unit that of the FCC catalyst), the difference in total acidity (in the 30–50 mmol gNH31 interval) and the average acidic strength (78–100 kJ mol1) are relatively small. However, there is a strong impact of these properties on the results, as will be discussed later on this manuscript. In addition, the crystallinity of the catalysts has been analyzed through high-angle X-ray diffraction (XRD) techniques in a Philips XPert PRO automatic diffractometer operating at 40 kV and 40 mA, in the u –2u configuration (Bragg Brentano geometry), with Cu Kα (l = 1.5418 Å). Data were collected from 3.5 to 60 (step size = 0.02, time per step = 100 s, total time =8 h). The observed spectra are characteristic and well-established in the literature for the HZSM-5 [57,58], HBeta [59] and HY [60,61] zeolites. Reaction equipment and product analysis The catalytic cracking runs have been carried out in a CREC (Chemical Reactor Engineering Centre) riser simulator designed to operate in laboratory scale under similar conditions than FCC unit reactor [62]. A detailed schematization of the FCC simulator is provided in Scheme 1 [21]. The operating conditions, resembling those typically used in industrial units, were: temperature, 550 C; contact time, 6 s; catalyst/bio-oil mass ratio (in a dry base), 1.8–7.3 gcatgbio-oil1. Please note that, in our experimental approach, the
Table 2 Properties of the fresh zeolite-based catalysts. Physical properties
CY
CZ
CB
SBET (m2 g1) Micropore surface (m2 g1) Pore volume (cm3 g1) Micropore volume (cm3 g1) dp (Å) Unit cell size (Å) Chemical properties Zeolite percentage (wt %) Total acidity (mmol g1) Average acid strength (kJ mol1) B/L ratio (mol mol1)
122 96 0.15 0.04 117.3 24.30
143 36 0.24 0.02 62.2 –
133 32 0.32 0.02 89.6 –
15 30 100 0.75
20 50 78 0.22
20 40 85 0.24
contact time simulates the time that, in an industrial FCC unit, the catalyst spends inside the riser section of the FCC unit, from the bottom entrance to the exit towards the regeneration stripping section. A fixed raw bio-oil mass of 0.11 g was upgraded on each experiment, and the employed catalyst mass varied between 0.13 and 0.53 g. Blank runs in no catalyst conditions (using only sand and an α-alumina + bentonite mixture, respectively) allowed for gathering insights into the thermal effect contribution and corroborate the inert nature of the matrix and binder used in catalyst preparation. The product stream was extracted from the reaction chamber once the reaction time was completed through a rapidly triggered vacuum pump connected to an auxiliary chamber (30 cm3), and then analyzed on-line by means of gas chromatography (GC) in an Agilent Technologies 7890 A chromatograph provided with flame ionization (FID) and pulsed flame photometric detectors (PFPD). In a pre-experimental phase, MS analyses were used to accurately identify the reaction products, while for routine product analysis GC techniques were only applied. The products were grouped in lumps, as follows: CO, CO2, dry gases (C1-C2), liquefied petroleum gases (LPG, C3-C4), gasoline (C5-C12), light cycle oil (LCO, C13-C20) and coke. The amount of coke was quantified from the mass loss after combustion of spent catalysts in a TGA-Q 5000 TA Instruments thermobalance, following a heating rate of 3 C min1 from 300 C up to 550 C. The rest of the volatile products was quantified through analysis of a gaseous sample in a Varian CP-4900 Micro GC, provided with two modules equipped with thermal conductivity detectors (TCD): (i) a molecular sieve (MS-5, 10 m) where H2, O2, N2, CH4 and CO are separated and (ii) a Porapak PPQ (10 m) where H2O, CO2 and light products (methane, ethylene, ethane, propylene, propane, acetaldehyde, butanes, butenes and acroleyn) are separated.
Scheme 1. Detailed schematization of the FCC unit riser simulator [21].
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The yield of each product fraction (Yi) was defined as specified by Eq. 1: Yi ¼
mi 100 mbiooil
ð1Þ
where mi, is the produced mass fraction of product i; and mbio-oil corresponds to the mass of bio-oil (wet basis) o the mass of oxygenates in bio-oil (dry basis). The conversion of each oxygenate component (Xi) of bio-oil was calculated from Eq. 2: Xi ¼
ðmi Þt¼0 ðmi Þt 100 ðmi Þt¼0
ð2Þ
where (mi)t=0 is the mass content of each oxygenate in the raw biooil feed and (mi)t the corresponding amount in the remaining biooil. Different parameters have been determined for evaluating the quality of the produced liquid fuel [11]. The deoxygenation degree (DOD, in wt%) was calculated from the oxygen mass content in the liquid fuel, (mO)fuel, and the oxygenates in bio-oil, (mO)feed, as: ðmO Þf uel DOD ¼ 1 ð3Þ 100 ðmO Þbiooil The liquid fuel yield (Yfuel, wt%) was defined as: Y f uel ¼
mf uel 100 mbiooil
ð4Þ
where mfuel is the mass of liquid fuel (hydrocarbons C5+ and oxygenates in the same boiling point range) and mbio-oil is the mass of bio-oil. On the other hand, the oxygen/carbon mass ratio (Ox/C, gOgC1) was determined from the oxygen mass content in the fuel product, (mO)fuel, divided by the carbon mass content in the fuel product, (mC)fuel, as: Ox ðmO Þf uel ¼ 100 C ðmC Þf uel
ð5Þ
Results and discussion
Fig. 1. Evolution with the catalyst/bio-oil mass ratio of the yields (on a wet basis) of (a) carbon products, (b) water and (c) coke, for the different catalysts.
Yields and product distribution The total yields of the main cracking products resulting from the dehydration of bio-oil (carbon volatile products and water) together with the yield of carbonaceous material deposited on the catalyst (coke) are displayed in Fig. 1. This Figure also displays the results of the thermal cracking (in a sand bed and without a catalyst) at C/O = 0, as well as the original carbon product and water content of the raw-bio oil. For the three studied catalysts, the yield of volatile carbon products (Fig. 1a) decreases upon increasing the catalyst/bio-oil ratio, while the yields of water and coke increase (Fig. 1b and c, respectively), as a direct consequence of higher catalyst/bio-oil ratios enhancing deoxygenation reactions (dehydration, decarbonylation and decarboxylation) when a higher amount of active catalyst sites become available. In addition, the C/ O ratio has a notable effect over coke yields because acid catalyst sites activate their formation mechanisms [63,64]. The results in Fig. 1 indicate that the reactions implied in bio-oil cracking are very rapid at 550 C, and thus become highly evolved even at the lowest catalyst/bio-oil ratio (1.8 gcatgbio-oil1), in such a way that, upon a further increase of the catalyst/bio-oil ratio, only slight variations are observed on dehydration and coke formation reactions. The yields of each product fraction in Fig. 1 are dependent on both the utilization of a catalyst and also the catalyst properties (mainly shape selectivity and secondarily catalyst acidity). This
way, the water yield trend (CZ > CB > CY, Fig. 1b) is in agreement with the total acidity trend (see Table 2) of the catalysts, known to enhance deoxygenation reactions [65]. Coke yields are overall very low in contrast to carbon volatile products (0–4 wt% vs 52–55 wt%, respectively). As observed in Fig. 1c, the CY catalyst is the one that yields the highest amount of coke (2.4–4.0 wt%). This can be attributable to its porous structure, containing cages formed in the intersections between the crystalline channels (which are not present in the structure of the HBeta and HZSM-5 zeolites), within of which coke precursors are trapped and further evolve towards polyaromatic structures, reaching similar coke yields to those obtained in the cracking of conventional FCC feeds [26]. It should be highlighted that, despite catalyst coking taking place, no operational drawbacks were observed upon feeding the bio-oil into the FCC simulator device. The fast injection in the setup, together with the significant water content in the bio-oil (known to prevent coke formation) and the severe temperature used in the process favor the fast vaporization of the feed and hence its fast conversion into desired products. On the other hand, the catalysts prepared from HZSM-5 and HBeta zeolites present more severe shape selectivity with narrower pores (see Table 2) which partially limit the formation of polyaromatics within the channels of the zeolite. Despite the well-known capacity of the HZSM-5 zeolite for minimizing coke
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formation, due to its microporous structure with channels that limit the progress of bimolecular hydrogen transfer and aromatic condensation reactions [53,66], the coke yields of the CZ catalyst (0.8–1.8 wt%) are overall higher than those of the CB catalyst (0.4– 1.0 wt%), which is attributable to the higher acidity of the CZ catalyst also playing a key role in coke condensation reactions [67]. The differences in lump product yields (in a dry basis) at different C/O ratios for the three catalysts are displayed in Fig. 2. The results at C/O = 0 correspond to a blank reaction (no catalyst and using sand). CO and CO2 formation reactions are favored upon increasing the catalyst/bio-oil ratio for all the catalysts (19.2– 27.6 wt%, Fig. 2a). CO and CO2 yields are higher for the CZ catalyst (22.5–27.6 wt%) as a consequence of decarboxylation and decarbonylation reactions being enhanced by catalyst acidity. On the other hand, the performance of the CY and CB catalysts resulted very similar (19.2–24.4 wt%) and, while the pore structure of the former (with wider average pores) enhances the diffusion for the oxygenates in bio-oil, the latter possesses a higher acidity that can outperform catalysts with overall wider pores. The compromise of the CB catalyst, in terms of oxygenate accessibility to the acidic sites (medium shape selectivity severity), and acidic sites themselves (intermediate acidity, Table 2), justifies its similar dry gas (14.8–19.5 wt%, Fig. 2b) and LPG (11.3–13.2 wt%, Fig. 2c) yields to those of the CZ catalyst (14.8–17.6 wt% and 12.1– 14.5 wt% for dry gas and LPG, respectively). On the other hand, these yields are lower for the CY catalyst (10.3–12.6 wt% dry gas
and 6.8–7.5 wt% LPG) due to its lower acidity and presumably more rapidly deactivated according to the results for coke yields (Fig. 1c, Fig. 2f). The gasoline yield (Fig. 2d), which resulted the main cracking product (34.0–55.3 wt%), follows the CY > CB > CZ trend, in opposite order to that of the total catalyst acidity. The lower acidity of the CY catalyst seems to be key for minimizing gasoline overcracking reactions towards dry gas and LPG, thus enhancing the yields of liquid fractions, which are also favored by operating at lower catalyst/bio-oil ratios. The velocity at which the transformation of bio-oil takes places is evidenced from the very low yields of LCO that were attained with all the catalysts (0.7–1.3 wt%, Fig. 2e), evidencing that the conversion of this fraction into lighter products is immediate, and negligible differences were observed between the three catalysts. This indicates that the different microporous structure of the corresponding zeolites does not have a strong impact in terms of diffusional limitations of LCO components. As discussed earlier on the results in Fig. 1, the higher coke yield (4.2–7.0 wt%, Fig. 2f) is attained with the CY catalyst due to the different porous structure of zeolite HY in comparison with the other two catalysts, lacking macropores in their matrix and without cages in their pore intersections, hence favoring the diffusion of coke precursors towards the outside of the catalyst. The lower coke yield is for the CB catalyst (0.7–1.7 wt%) thanks to a compromise between moderate shape selectivity and acidity in comparison with the other catalysts.
Fig. 2. Evolution with the catalyst/bio-oil mass ratio of the product yields (in a dry basis) for the different catalysts.
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In the catalytic cracking of bio-oil, considerable amounts of hydrocarbons are produced, from complex combinations of cracking, deoxygenation, dehydration, decarboxylation, decarbonylation, and alkylation reactions, among others. Fig. 3 shows the total hydrocarbon yields (in a dry basis) (Fig. 3a), and also those corresponding to the liquid C5-C20 fractions (Fig. 3b) that can be potentially blended in the gasoline and diesel pools. The higher total hydrocarbon yield corresponds to the CB catalyst due to the aforementioned compromise between its shape selectivity and acidity, being of 42.6 wt% at a catalyst/bio-oil ratio of 7.3 gcatgbio1 oil , and slightly higher than that corresponding to the CZ catalyst (40.2 wt%) at the same conditions. Regarding the C5-C20 fraction (Fig. 3b), the higher yields attained with the CB catalyst are clear, reaching a maximum value of ca. 14 wt% for a catalyst/bio-oil ratio in the 4–5 gcatgbio-oil1 range. At higher C/O values, overcracking reactions acquire greater relevance and, consequently, liquid hydrocarbon fractions are lower. Table 3 allows for a detailed evaluation of the effect of the catalyst/bio-oil ratio and the different catalysts on the yield of the dry gas (C1-C2) and LPG (C3-C4) fractions. The C1-C2 (dry gas) yields are overall lower for the CY catalyst, along the same lines of the yields reported in Fig. 2b and d for this catalyst. The comparison between the results obtained with the CZ and CB catalysts is intricate and, even though a lower methane yield (7.7–8.5 wt%) was attained with the CZ catalyst, the rest of the yields do not present significant differences in the studied C/O range. However, it was observed that for a catalyst/bio-oil ratio of 7.3 gcatgbio-oil1, the CB catalyst provides a higher ethylene yield (6.7 wt%), which might be of industrial interest, at the expense of a higher methane production. Regarding the LPG fraction (C3-C4 range), operating with the CY catalyst leads mainly to olefins like propylene (3.0– 4.3 wt%) and butenes (1.9–2.2 wt%). Notably, the yields of butanes are higher with the CB (1.1–2.4 wt%) and, particularly, with the CZ catalyst (2.7–4.2 wt%). An overall decreasing tendency was observed in terms of the olefinic nature of the individual C2, C3
Fig. 3. Evolution with the catalyst/bio-oil mass ratio of (a) the total hydrocarbon yield (in a dry basis) and (b) C5-C20 hydrocarbon yields for the different catalysts.
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Table 3 Gas product yields (wt%) at the lowest and highest studied catalyst/bio-oil ratios in the cracking of raw bio-oil over different zeolites.
Methane, CH4 Ethylene, C2H4 Ethane, C2H6 Propylene, C3H6 Propane, C3H8 Butenes, C4H8 Butanes, C4H10
Catalyst/bio-oil (gcatgbio-oil1)
CY
CZ
CB
1.8 7.3 1.8 7.3 1.8 7.3 1.8 7.3 1.8 7.3 1.8 7.3 1.8 7.3
5.7 6.9 3.0 3.6 1.6 2.0 4.3 3.1 0.2 1.0 2.2 1.9 0.1 1.6
7.7 8.5 4.9 5.9 2.0 3.0 5.6 4.8 0.8 0.9 3.1 2.7 2.7 4.2
8.2 9.7 4.7 6.7 2.1 3.1 5.1 5.8 0.9 1.3 4.2 3.7 1.1 2.4
and C4 fractions upon increasing the catalyst/bio-oil ratio as a consequence of hydrogen transfer reactions being favored due to the higher amount of available acidic catalyst sites. No clear differences are observed among the catalysts for the transformation of ethylene into ethane. On the other hand, in the propylene to propane transformation, and even more in that of butenes into butanes, the different catalyst acidity seems to acquire importance, explaining the CY > CB > CZ olefinity trend of the C4 fraction, with an opposite order to the acidity trend. Conversion of individual bio-oil oxygenates The effect of the catalyst/bio-oil ratio on the conversion of different individual oxygenate compounds (representative of the different chemical families in bio-oil) for the three catalysts is displayed in Fig. 4. These compounds are: acetaldehyde (from lineal aldehydes), furfural (from furans), acetic acid (from acids), 3methyl-1,2,cyclopentenone (from cyclic ketones), hydroxyacetone (from lineal ketones), phenol (from branched phenolics), levoglucosan (from high molecular weight sugars), and phenolic compounds (to also consider the results corresponding to phenolic ethers, like guaiacol, syringol, and their derivates). The catalytic cracking mechanisms for biomass-derived oxygenates is wellestablished and, as reported by Huber and Corma [1], hydrogen producing and consuming reactions are of critical importance. The conversions of acetaldehyde (Fig. 4a), hydroxyacetone (Fig. 4e) and levoglucosan (Fig. 4h) are considerably high even for a catalyst/bio-oil ratio of 1.8 gcatgbio-oil1 for all the catalysts. The high acetaldehyde conversion might imply its partial undesired transformation into coke, given the important role that this component plays in the reactions involved in the formation of pyrolytic lignin due to the formation of trimethyltrioxane as intermediate [29]. On the other hand, the high conversion of levoglucosan (present in high proportion in bio-oil derived from fast pyrolysis) can be explained from its high instability and reactivity [68]. On the contrary, the conversion of furfural (Fig. 4b) and 3-methyl-1,2-cyclopentenone (Fig. 4d) is lower at the studied conditions, increasing up to almost full conversion (94.7 wt%) with the most active CZ catalyst for a catalyst/bio-oil of 7.3 gcatgbio-oil1. The conversion for acetic acid (Fig. 4c), phenol (Fig. 4f) and the phenolic groups in bio-oil (Fig. 4g) is lower and, even if it increases upon increasing the catalyst/bio-oil ratio, it is still far from reaching full conversion, with acetic acid being the least reactive component. Horne and Williams [69], reporting studies of furfural cracking over HZSM-5 zeolites, mainly obtained aromatic products (mostly alkylbenzenes) due to the higher acidity of their catalyst and the high space time, which facilitates olefin condensation. Bertero et al. [32]. using a fixed bed MAT reactor at
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Fig. 4. Evolution with the catalyst/bio-oil ratio of individual conversions of the different oxygenated families in the cracking of raw bio-oil over different catalysts.
500 C and 60 s of time on stream, and an equilibrated FCC catalyst, produced majoritarily pentenes from the cracking of furfural and aromatics from the cracking of 3-methyl-1,2-cyclopentenone. The reactivity order of the bio-oil oxygenates in the reaction media (corresponding to raw bio-oil), is in agreements with the results reported in the literature for the cracking of these components individually feed (at different conditions and using different catalysts) [30,29,22,70,71]. When comparing the results for the different catalysts, the conversion order, with the exception of phenol and phenolic compounds, is the following: CZ > CB > CY, which corresponds to the acidity trend of the catalysts. These results indicate that the total acidity of the used catalyst will have a greater impact on the conversion of individual oxygenates than their pore network structure.
This situation also takes place in the conversion of levoglucosan, a voluminous unstable molecule, whose conversion is significant at the studied temperature even without a catalyst [71], and almost complete at low catalyst/bio-oil ratios (Fig. 4h). However, the conversion of levoglucosan over the CZ catalyst is only slightly higher than over the CB catalyst, which seems to indicate that the HZSM-5 zeolite presents diffusion limitations (narrower average pores, see Table 2) for the levoglucosan molecule, which is favored by HBeta zeolite. In the conversion of phenol and phenolic groups, the conversion order does not follow the increasing acidity trend, and the CB catalyst appears as the most active, which points out a compromise between catalyst acidity and pore accessibility to acidic sites in the case of this catalyst. The low activity of the CY catalyst should be highlighted in the conversion of these compounds, attributable to
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its lower acidity in fresh conditions as well as the fact that it might already be considerable deactivated at such short reaction time as 6 s, based on a previous work by Ibarra et al. [21].
Table 4 Quality indexes of the liquid cracking product of raw bio-oil obtained with the different catalysts at the lowest and highest studied catalyst/bio-oil ratios. Parameters
catalyst/bio-oil (gcatgbio-oil1)
CY
CZ
CB
DOD
1.8 7.3 1.8 7.3 1.8 7.3
61.1 70.7 33.1 27.2 0.50 0.38
70.7 80.4 26.7 19.7 0.38 0.31
67.7 77.3 29.8 22.1 0.40 0.34
Quality of the fuel Table 4 summarizes the indexes for assessing the quality of the liquid product from raw bio-oil cracking: deoxygenation degree (DOD, Eq. 3), liquid fuel yield (Yfuel, Eq. 4), and oxygen/carbon ratio (Ox/C, Eq. 5). The extent of the DOD follows the CZ > CB > CY trend, which highlights the fact that catalyst acidity features will not only heavily determine product distribution, but is also a key property for increasing oxygen removal by enhancing cracking and deoxygenation (dehydration, decarboxylation and decarbonylation) reactions of the most part of bio-oil components. All in all, the deoxygenation efficiency is high for the three catalysts (61.1–80.4 %) increasing with the catalyst/bio-oil ratio (higher availability of acidic sites for reactants). Specifically, using the CZ catalyst, DOD indexes from 70.7% to 80.4% were attained. However, a higher oxygen removal in the products will come at the expense of lower liquid fuel yields (Yfuel, 22.1–33.1 wt% in the studied C/O range) and higher coke and gas product yields (see results in Fig. 2) and, therefore, a compromise between the different operation parameters (process severity) and desired production goals must be established. An increase in the catalyst/bio-oil ratio also shows a tendency for slightly decreasing the Ox/C ratio (defined in Eq. 5) in the liquid product, which is a key quality index to evaluate the fuel. Upon decreasing the Ox/C ratio, the H/C ratio in the fuel is higher (from
379
Yliquid Ox/C
fuel
0.2 to 0.4 in the catalyst/bio-oil range of 1.8–7.3 gcatgbio-oil1 for the CY catalyst). It was observed that the lowest Ox/C value was attained at a catalyst/bio-oil ratio of 7.3 gcatgbio-oil1 and using the CY catalyst which, with its higher acidity, provides a high-quality liquid fuel with a reduced oxygen content (80.8% removal). However, as mentioned earlier, this also implies a lower liquid product yield (27.2 wt%) in contrast to lower catalyst/bio-oil ratios, due to excessive overcracking of the gasoline fraction towards dry gas and LPG. Yields of the gasoline fraction components The gasoline fraction consists of C5-C12 hydrocarbons and oxygenates in the same boiling point range and, therefore, assessing both fractions separately is advisable. Fig. 5 displays the yields corresponding to the main oxygenate and hydrocarbon families in the gasoline fraction for the different catalysts, where important effects of both the catalyst and catalyst/bio-oil ratio
Fig. 5. Comparison of the yields of hydrocarbons (graph a and b) and oxygenates (graph c and d) in the gasoline fraction obtained in the cracking of raw bio-oil over different catalysts at the lowest and highest studied catalyst/bio-oil ratios.
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were observed. For a catalyst/bio-oil ratio of 1.8 gcatgbio-oil1, overall, the main hydrocarbons are olefins (3.1–4.3 wt%), with lower yields of i-paraffins (2.1–3.3 wt%), aromatics (1.6–3.3 wt%), naphthenes (1.3–2.8 wt%) and n-paraffins (0.3–0.7 wt%), as seen in Fig. 5a, attaining total hydrocarbon yields of 10.1–12.2 wt%. An increase in the catalyst/bio-oil ratio up to 7.3 gcatgbio-oil1 (Fig. 5b) notably modifies the hydrocarbon composition in gasoline, with significantly lower olefin and naphthene yields. The former are mainly affected by enhanced condensation of aromatics (which increase up to 4.8 wt%), while the latter mostly undergo cracking reactions towards paraffins and olefins, which later on will likely also condense into aromatics. Interestingly, total hydrocarbon yields (11.9–12.3 wt%) are very similar to those obtained at lower catalyst/bio-oil ratio. At this high catalyst/bio-oil ratio the mildly acidic CB catalyst presents more advantages, with i-paraffins being the main chemical group (5 wt%), while the aromatic yield only increases slightly in comparison with the lowest catalyst/bio-oil value, accounting for only 50% of the yields attained with the other catalysts (2.3 wt% vs 4.7–4.8 wt%). Furthermore, the yield of olefins, even if higher than for the other catalysts, is still notably lower than at lower catalyst/bio-oil value. The higher acidity of the CZ catalyst in contrast to the CY catalyst had an enhancing effect over cracking and deoxygenation reactions, providing higher yields of hydrocarbons (with the exception of aromatic compounds). However, at the same time, the transformation of phenol and phenolic compounds is hindered within the pore structure of the HZSM-5 zeolite due to steric difussional limitations. Likewise, these limitations restrict Diels-Alder condensation reactions of olefins towards
aromatic compounds and, according to our results, also cracking reactions of naphthenes. It should also be considered that, since olefins are primary products, their formation is favored by catalyst acidity, which at the same time also enhances their transformation into n-paraffins and, in a lesser extent, into iparaffins. On the other hand, the better results achieved with the CB catalyst, among which lower aromatic and higher naphthene yields outstand, are attributable to a good compromise between moderate acidity and shape selectivity, which not only hinders olefin condensation into aromatics, but also olefin cracking (into dry gas and LPG) and naphthene cracking. On the other hand, the yields of oxygenates in the gasoline fraction are complementary results to those of the conversion of individual components in bio-oil, discussed in Fig. 4. As seen from Fig. 5c and Fig. 5d, oxygenated compounds constitute the main fraction in gasoline with yields in the 34.6–45.0 wt% and 21.9– 33.2 wt% range, at 1.8 and 7.3 gcatgbio-oil1, respectively. Upon increasing the catalyst/bio-oil ratio, gasoline becomes less oxygenated with deoxygenation being favored, as previously discussed from the results in Table 4, due to a higher availability of active sites. Additionally, it was observed that the least acidic CY catalyst provides higher oxygenate yields (with the exception of ketones), due to the lower activity of the CY catalyst in cracking and deoxygenation reactions. The yields of oxygenates decrease with more acidic catalysts. The lower yield of acetaldehyde and phenol (with the CZ and CB catalysts respectively) should be highlighted, since these results are relevant for minimizing issues derived from the condensation of phenolic components in bio-oil storage and handling.
Fig. 6. Comparison for different catalysts of the distribution of the yields of (a) aromatics, (b) n-paraffins, (c) olefins and (d) i-paraffins with different C atom number in the gasoline fraction. (Catalyst/bio-oil ratio = 1.8 gcatgbio-oil1).
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Delving into the results in Fig. 5, interesting differences were observed upon evaluating individual yields of each hydrocarbon family (on a carbon number basis) in the gasoline fraction, as seen in Fig. 6. At a low catalyst/bio-oil ratio (1.8 gcatgbio-oil1), the aromatic components detected are mainly in the C6-C8 range (Fig. 6a), and with notably lower yields for the CB catalyst. With respect to the comparison between the CY and CZ catalysts, steric hindrances are again evidenced, with the microporous HZSM-5 structure limiting the formation of aromatics, either through cracking of phenolics or olefin condensation. The yield of n-paraffins (Fig. 6b) is very low (<0.2 wt%), and therefore the measurements are liable to a higher experimental error. A heterogeneous distribution of the n-paraffins is observed in terms of carbon atoms, with both light and heavy paraffins being present, and which is also affected by catalyst properties. Olefins (Fig. 6c) are mainly C5-C6, with higher yields for the CB catalyst which, as previously discussed, has a limited capacity for transforming these olefins (either through aromatic condensation of light olefin cracking reactions). In the same way, C5 and C6 iparaffins are predominant (Fig. 6d), with higher yields for more acidic catalysts (CB, CZ), which indicates that the higher acidity is enhancing their transformation from olefins through hydrogen transfer reactions. Conclusions The shape selectivity and acidity of the zeolite in the catalyst have a significant impact on catalyst activity, selectivity and coke deposition on the cracking of raw bio-oil in FCC conditions. The commercial FCC (HY zeolite-based) catalyst is adequate for maximizing liquid fuel yields, with a higher gasoline yield and a limited gas product formation. At the same time, coke yield is favored in this catalyst due to the higher retention capacity of coke precursors. On the other hand, the more acidic HZSM-5-based catalyst is suitable for producing light olefins and the HBeta-based one is intermediate between the other two. The concentration of hydrocarbons in the gasoline fraction, with olefins and aromatics as the main compounds, is dependent on the porous structure and acidity of the zeolites. Consequently, the aromatic content is higher in the gasoline produced using the HY zeolite and lower with the HZSM-5 zeolite. Furthermore, the cracking results are dependant on the composition of bio-oil, due to the distinct reactivity of its components, in the following order: sugars (levoglucosan) >lineal aldehydes (acetaldehyde) >lineal ketones (hydroxyacetone) >furans (furfural) cyclic ketones (3-methyl-1,2-cyclopentenone) >acids (acetic acid) branched phenols (phenol) phenolic groups. Acknowledgements This work has been carried out with financial support from the Ministry of Economy and Competitiveness of the Spanish Government (MINECO) (Projects CTQ2012-35192, co-funded with FEDER funds) and the Basque Government (Project IT748-13). Dr. Idoia Hita is grateful for her postdoctoral grant awarded by the Department of Education, University and Research of the Basque Government (grant number POS_2015_1_0035). Dr. Alvaro Ibarra is grateful for his PhD grant from the University of the Basque Country (UPV/EHU). Authors also acknowledge Ikerlan/IK-4 and Petronor S.A. for providing the raw bio-oil and the catalyst used in this work, respectively. References [1] G.W. Huber, A. Corma, Angew. Chemie Int. Ed. 46 (38) (2007) 7184. [2] S. Gunukula, S.J.W. Klein, H.P. Pendse, W.J. DeSisto, M.C. Wheeler, Appl. Energy 214 (2018) 16.
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