Ketal Sugar Conversion Into Green Hydrocarbons by Faujasite Zeolite in a Typical Catalytic Cracking Process.

Fluidized catalytic cracking (FCC) converts hydrocarbons in the presence of a catalyst based on faujasite zeolite (USY and REY). While hydrocarbon is poorly reactive, biomass and its derived compounds are highly functionalized and not suitable to a typical FCC process. To overcome this limitation biomass was first converted into a dense and stable bio-crude composed mainly of ketal-sugar derivatives by using acetone in diluted acid. Here, a representative compound of this bio-crude, 1,2:3,5-di-O-isopropylidene-α-D-xylofuranose (DX) in n-hexane, was converted by USY and a commercial FCC catalyst containing USY, at 500°C, in a fixed bed and fluidized bed reactors, respectively. Faujasite Y is very efficient in converting DX. More than 95% conversion was observed in all tests. Over 60 wt.% was liquid products, followed by gas products and only around 10% or less in coke. The higher the catalyst activity the greater the aromatics in the liquid products and yet higher coke yields were observed. In particular, simulating more practical application conditions: using deactivated catalyst in a fluidized bed reactor, improved green hydrocarbons production (mono-aromatic up to 10 carbons and light hydrocarbon up to eight carbons) and unprecedented lower coke yield (≈5 wt.%) for bio-feeds. The present results further suggest that catalyst will play a primary role to convert the bio-crude into target hydrocarbons and overcome the transition of a non-renewable to a renewable refinery feed.

Introduction

Energy needs have been increasing substantially since the industrial revolution due to population growth and goods demand. Today oil, natural gas, and coal are the main sources for chemicals, fuel and energy production (Smil, 2004; Conti et al., 2016). But the unrestrained exploitation of these non-renewable resources almost doubled the greenhouse gas concentration in the atmosphere in the last 100 years (Russ and Criqui, 2007; North and Styring, 2015). Among several alternatives to reduce the carbon footprint, the production of green-fuels for typical refinery process can shorten our transition of a non- to a renewable refinery. As a huge ready-to-use structure for fuel products and distribution is available, and also requires simple modifications to the fuel-legislation and no adaptation in motors. Besides, simultaneously this avoids building an entire new structure (Goldthau, 2017). Thus, the conversion of biomass into regular fuel using a typical refinery is one of the most important issues that chemistry now faces (Ragauskas et al., 2006).

A major process of a typical refinery is fluidized catalytic cracking (FCC). This process produces gasoline, liquefied petroleum gas (LPG), light olefins and light cycle oil. The FCC process consists of three distinct steps: reaction, separation, and regeneration. The reaction occurs in the riser reactor in a few seconds under a reducing atmosphere due to the presence of hydrocarbons and hydrogen. After this step, catalyst and reaction products are separated by means of cyclones and a stripper. Finally, the spent catalyst with coke content in the range of 0.6 to 2.0 wt.% is directed to the regenerator vessel. The catalyst regeneration occurs at high temperature (above 700°C) and in the presence of O2, H2O, and other gases (O'Connor and Pouwels, 1994; Wen et al., 2002). The presence of steam can remarkably affect the catalyst properties (Escobar et al., 2005; Cerqueira et al., 2008) and a central point to develop laboratory test more close to typical FCC condition is to use deactivated or equilibrium catalysts. Also, the reaction-regeneration feature of the FCC makes this process eligible to co-process oil and renewable feeds (O'Connor, 2007).

The arrival of first generation of biofuels, such as bio-ethanol and bio-diesel allowed a partial replacement of fossil resources, also mitigate CO2 emissions. These new generation transport fuels are now facing key environmental and strategic questions, such as the fact that, they are derived from agricultural commodities that compete with food crops. In contrast, lignocellulosic biomass of forest and industrial waste are an attractive raw material to produce of renewable fuels (Ladanai and Vinterbäck, 2009; Langholtz et al., 2016). Biomass comprising cellulose, hemicellulose and lignin are commonly referred to as second generation biomass of and typically derived from non-edible residues such as bagasse and others from food industries. However, biomasses have low density and a broad diversity in functionalized components compared to hydrocarbons. Thus, it is mandatory to improve density and adjust the reactivity of any bio-feed prior to its conversion in a typical refinery process.

There are some strategies to introduce second-generation biomass in refineries (Huber et al., 2006; Serrano-Ruiz and Dumesic, 2011; Jong and Ommen, 2015). Pyrolysis and fast pyrolysis (Lappas et al., 2002; Oasmaa et al., 2003; Oasmaa and Meier, 2005; Mendes et al., 2016) have generally been accepted as primary processes for this transformation for further conversion into refinery (Adjaye and Bakhshi, 1995; Samolada et al., 1998). Pyrolysis can be performed in the presence of a catalyst (CPO) (French and Czernik, 2010; Graça et al., 2011), or post-treated by a hydro-deoxygenation (HDO) process (Saidi et al., 2014; Talmadge et al., 2014). The oils produced by these processes are typically dark, dense, viscous and composed of a complex mixture of different oxygen-containing molecules in function of the pyrolysis process, like phenols, sugars, carboxylic acids and 15–30 wt.% of water (Sfetsas et al., 2011; Silva et al., 2016). Thus, bio-oil is non-miscible with hydrocarbon, highly acid and corrosive. These bio-oils result in increase coke during co-processing with conventional feeds in FCC (Pereira and Benoit, 2013; Talmadge et al., 2014) and hence could be co-processed in very low concentration (Thegarid et al., 2014; Pinho et al., 2015). For instance, even after a severe hydrotreatment (H/C in bio-oil ≈1), high amounts of coke and gas was recorded (Table 1). The pyrolysis process solves the density, but not the reactivity problem of a bio-feed for further co-conversion at typical refinery conditions. However, bio-oil can be transformed into products by a coupling process, like hydrogen and synthesis gas (Trane et al., 2012) or fuels (Bulushev and Ross, 2011).

Table 1

Yield of oil products, gas, and coke used to convert bio-oil and oxygenated model compounds.

Bio feed, catalysts, temp.	Liquid wt.%	Carbon* wt.%	Gas wt.%	References
Raw bio-oil, ZSM-5, 490oC	7	19	44	Vitolo et al., 1999, 2001
20% hydro-treated bio-oil/80% gasoil, HY 500oC	40	6–12	20–40	Samolada et al., 1998
Raw bio-oil	26	32	–	Vispute et al., 2010
Hydro-treated Wood Bio-oil**	80***	13	22	Vispute et al., 2010

Include coke, char, and Tar.

100 bar of H.

Several classes of compounds.

Recently, we proposed an alternative approach to fuel production from second generation biomass in two steps (Batalha et al., 2014). The biomass was firstly converted in mild conditions into a bio-crude (density of 1.1 gmL−1 and CHO composition ≈ 58, 7, and 35% by weight, respectively) combining acid catalyzed hydrolysis with organic reactions such as ketalization (Garrett et al., 2015; de Souza et al., 2017) and acetylation (Durange et al., 2015). The biocrude is composed by a mixture of isopropylidene ketals containing mono and polyshacharides-ketals (Garrett et al., 2015). For instance 1,2:3,5-di-O-isopropylidene-α-D-xylofuranose (DX) and 1,2: 5,6- Isopropylidene- α-D -glucofuranose (DG) were produced in high concentration, 50 wt.%. A later publication also applied similar condition to depolymerization wood biomass and avoided undesirable reaction of sugars (Questell-Santiago et al., 2018).

Secondly, the feasibility of producing hydrocarbons, especially aromatics, using typical model components of the ketal-bio-crude such as DX and DG were shown by a simplified catalytic protocol (Batalha et al., 2014, 2016). Our approach of circular economy for producing green-aromatics can be illustrated in Figure 1.

Figure 1

Circular economy based on second generation biomass converted into green-aromatics and -light hydrocarbons in three steps: (A) biomass production, sugarcane bagasse: the residue of ethanol and sugar industry, as feed; (B) biomass conversion into biocrude: sugar-ketal, ex. DX, DG, and polysaccharides-ketal (Garrett et al., 2015); (C) biocrude conversion into hydrocarbons.

The present work is focused on the catalyst and the catalytic testing conditions, using Y zeolite, the workhorse, active component of the FCC catalyst. First, with a fixed bed reactor, using pure USY zeolites comprehensively characterized, we showed that the DX in hexane was efficiently converted and the transformation was sensitive to catalysts properties. For example, USY was covered by silica to demonstrate that the catalyst controls the DX conversion. Then, approaching the conditions in commercial practice, a mixture of 30 wt.% of DX in n-hexane were converted in the presence of a commercial FCC catalyst, fresh or deactivated, in a fluidized bed reactor. For each process an overall view of the material balance and product distribution in gas, liquid and coke fractions, are discussed. The products distribution and catalytic properties provided insights of DX conversion into hydrocarbons.

Experimental Section

Catalysts

The faujazite zeolite USY (provided by Petrobras) was thermal treated in a furnace and the temperature increased from room temperature to 500 at a rate of 10°C min−1. At the final temperature, deionized water was pumped at 1 mL min−1 for 1 h. The catalyst obtained through this procedure was named USY-D500. The commercial catalyst was provided by Fabrica Carioca de Catalisadores S.A. and it was used as provided and after thermal-treatment (at a final temperature of 720°C and following the same procedure of the USY used in the fixed bed test). An additional treatment process was carried out at the USY-D500: 4 g of catalyst was dispersed under magnetic stirring in 100 mL of n-hexane and heated to 70°C under reflux. When the mixture reached 70°C, 0.6 mL of tetraethylorthosilicate (TEOS) was added. After 1 h, the mixture was filtered and the sample dried at 100°C. After 24 h, the catalyst was calcinated at 500°C for 3 h. This procedure resulted in the catalyst namely USY-D500-Silica. Prior to use, the zeolite was submitted to ion exchange using a NH4NO3 2M solution, at 80°C, for 1 h. The ion exchange procedure was performed 4 times, followed by calcination at 450°C for 4 h under airflow. This procedure resulted in no sodium being detected in the final sample as previously reported in the literature.

Catalytic Cracking

The n-hexane cracking in differential reactor was carried out in a high throughput unit (Ferreira et al., 2014). All experiments were conducted at room pressure (open reactor) in a continuous down flow fixed-bed micro reactor. Prior to catalyst evaluation, all the catalyst were heated simultaneously under N2 flow from room temperature to the desired temperature (773 K), at a rate of 10 K/min. Then the reactor flow was modified to a mixture of n-Hexane 10% v/v in N2 (30 mL/min) for all catalytic tests. A typical run was carried out with 0.01 g of catalyst and reaction products were analyzed on-line after three distinct times on stream (3 min, 17 min and 32 min) by gas chromatography using a Shimadzu GC-2010 (Maia et al., 2010). The activities and selectivity were presented as the average result obtained after 17 min and 32 min on stream. The error in conversion and selectivity of n-hexane test was determined in previous working and are <5 and 2%, respectively (Ferreira et al., 2014).

Reactions performed in a fixed bed reactor (Figure S1) occurred at 500°C, under atmospheric pressure and nitrogen flow (100 mL min−1) and 400 mg of catalyst. Prior to reaction, the catalyst was heated under a nitrogen flow (100 mL min−1) from room temperature to 500°C at a rate of 10°C min−1, remaining at this final temperature for 30 min. During the reaction, a 0.2 mL min−1 liquid flow of DX in n-hexane, at room temperature, was directly to the reactor entrance and mixed with N2 (100 mL min−1). For comparison, pristine n-hexane cracking was carried out with the same reaction protocol and catalyst mass. The fluid catalytic cracking unit used to convert the new bio-feed in this study is presented in Figures S2, S3 (Pinto et al., 2017). The tests were carried out by using 10 mL of feed (pristine n-hexane and mixtures of DX of 30 wt.% in n-hexane) injected over 1 min, and 20 g of the fresh and deactivated catalyst were used in the catalytic bed. Before the reaction, the catalyst was activated in a nitrogen atmosphere for 12 h at 773 K. Nitrogen flow was calibrated at room temperature, and nitrogen flow at 200 mL min−1 (estimated at 500°C using the state equation of ideal gas) was used to fluidize the catalyst. The reactor was operated in between 450 and 500°C considering the height of the fluidized catalytic bed (Figure S3). OPAL is a fresh commercial FCC catalyst supplied for this work by Fábrica Carioca de Catalisadores S.A., and was used with particle diameter 115–200 mesh equivalent to a particle size diameter of 0.125–0.08 mm, respectively, but the composition is unknown.

The reaction products were distributed in: liquid, gas and coke. Three tests were carried out in the presence of the catalyst USY-D500 to provided an estimative of the error (Table S1). These values were 5, 4, and 6% for gas, liquid and coke, respectively. An error of 6% for all fractions was assumed as presented in Figure 5. The liquid fraction was obtained through condensation (−15°C in fixed bed and −18°C in the FCC test) by means of a condenser placed right after the reactor exit. The liquid amount was obtained directly by weight difference of the condenser before and after the reaction.

The liquid faction from the cracking reactions were analyzed off-line by both GCMS and CGFID. The determination of liquid products was obtained based on GCMS and the quantification was carried out in a GCFID. The GCMS system is an Agilent Technologies 7890A CG coupled to a 5975C MS in electron impact mode, an Agilent HP-5MS column was used and the oven temperature kept at 303 K for 7 min followed by a ramp to 443 K for 40 min, helium was used as carrier gas. At the inlet, a split ratio of 20:1, 14 psi pressure and 563 K was used. The GCFID system is an Agilent Technologies 7890A GC and the same method as GCMS was used.

The liquid products were obtained by subtracting the liquid mass by the mass of non-reacted n-hexane (determined according Text S1). The wt.% of each aromatic compounds in the liquid fraction was obtained multiplied the FID area (%), ARi, by the respective chromatographic factor (f) that is 0.82 for aromatic compounds. The wt.% of the remaining products were equal to each FID area (%) by assuming fi = 1. Remaining products, P, are determined by the sum of all olefins (containing five or more carbons), all hydrocarbons (containing five or more carbons), all light compounds (containing four or fewer carbons), heavy (assuming f = 1) and non-determined compounds (assuming f = 1). The total area of liquid products was determined by the sum of aromatics (corrected by each aromatic factor) and Pj (total of the remaining products), equation 1-a. The aromatic/total products ratio (SAR) was as presented in equation 1-b. These ratios are Sole (obtained by equation 1-c), Ssat (equation 1-d), and Slight (equation 1-e). Yields were obtained by multiplying the above ratios by the mass liquid products (obtained by discounting the amount of n-hexane, not reacted).

The gas products composition (H2, CO, CO2 methane, and hydrocarbon up to C4) were analyzed on-line using an Agilent Technologies Micro-GC 490. The amount of gas produced during the reaction was determined by the difference of water displaced during the reaction and the one in pure nitrogen flow (always quantified before reaction). The gas composition obtained was estimated using the average of 5 injections made on-line during time on stream of the reaction.

The amount of coke in the spent catalysts was determined through thermogravimetric analysis (Netzsch TG-IRIS). The samples were heated, under helium atmosphere, from 35 to 250°C at a rate of 10°C min−1 under N2 atmosphere. The temperature was kept at 250°C for 30 min, after which the atmosphere was changed to synthetic air (20.9% O2 in N2) and temperature increased to 700°C at a rate 10°C min−1 and afterwards it stayed on an isotherm for 30 min. The amount of coke in the catalyst corresponded to the weight lost at temperatures higher than 250°C and the coke yield was estimated comparing this value (considering the amount of catalyst in each reaction) to the total feed introduced into the reactor.

The 1,2:3,5-di-O-isopropylidene-α-D-xylofuranose (DX) was synthesized through the reaction of D-xylose (98.5%, VETEC, 20 g) with acetone (99%, VETEC, 400 mL) in the presence of sulfuric acid (P.A., VETEC, 16 mL). The reaction was carried out at 0°C and under mechanic stirring. The sulfuric acid was slowly added to the xylose/acetone suspension to avoid any changes to the mixture temperature. Once all acid was added, the mixture was stirred for 5 h at 0°C, and then a solution of NaOH (50 wt.%) was added until pH of 5. Complete neutralization was obtained with sodium bicarbonate (VETEC). After neutralization, the solution was filtered and all acetone removed under vacuum. The remaining liquid was dissolved in ethyl acetate (99%, VETEC) and the organic phase washed with distilled water. The solvent from the organic phase was removed under vacuum and the remaining liquid was dissolved in n-hexane (VETEC 99%). The part that was insoluble in n-hexane was discarded. N-hexane was then evaporated to obtain pure DX, which was weight (yield around 25%) and dissolved once more in n-hexane to obtain the mixtures used in the reactions (10 and 30 wt.% in n-hexane). This mixture was kept at 5°C to avoid n-hexane evaporation and DX degradation. The purity of the final compound was verified after each batch synthesis, through GCMS analysis.

Results and Discussion

Catalyst Characterization Used in the Fixed Bed Test

Fresh and modified USY by thermal treatment and silication are presented in Table 2. The thermal treatment in steam at 500°C slightly reduced both BET and external area compared to the USY precursor. The silication process of the USY-D500 (USY-D500-Silica) produced further decreases in the textural properties. Figure 2A shows typical type IV isotherms (IUPAC classification) of USY zeolite, with the occurrence of hysteresis cycle in the range of relative pressures 0.45