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DOI: 10.1039/C5RA22137D (Review Article) RSC Adv., 2015, 5, 103999-104042

Effective parameters on selective catalytic hydrodeoxygenation of phenolic compounds of pyrolysis bio-oil to high-value hydrocarbons

Hoda Shafaghat, Pouya Sirous Rezaei and Wan Mohd Ashri Wan Daud*
Department of Chemical Engineering, Faculty of Engineering, University of Malaya, 50603 Kuala Lumpur, Malaysia. E-mail: h.shafaghat@gmail.com; pouya.sr@gmail.com; ashri@um.edu.my; Fax: +60 3 79675319; Tel: +60 3 79675297

Received 22nd October 2015 , Accepted 27th November 2015

First published on 30th November 2015

Abstract

Pyrolysis bio-oil is recognized as a renewable and carbon-neutral fuel which could be a potential alternative for depleting fossil fuels. However, bio-oil is highly oxygenated and needs to be upgraded prior to be used as fuel additive. Catalytic hydrodeoxygenation (HDO) is an efficient technique for bio-oil upgrading. The reaction pathway for HDO of bio-oil is unknown since it is a mixture of hundreds of different compounds. The study on mechanism of transformation of these compounds could be helpful to propose an overall pathway for HDO of bio-oil. Phenols which are derived from pyrolysis of lignin fraction of biomass are considered as attractive model compounds for study of bio-oil HDO since they are highly stable in HDO reaction. Reaction pathway and product selectivity in HDO of phenols are highly affected by type of catalyst promoters and supports, catalyst preparation procedure, solvent type, chemicals used as co-feed and operating conditions (i.e., temperature and pressure). The effects of these factors on selective production of high-value hydrocarbons of aromatics and alicyclics from HDO of phenol, cresol, guaiacol and anisole are discussed in this review.

1. Introduction

Increase of energy request, depletion of fossil fuel resources as well as growing environmental concerns have raised the importance of finding renewable and environmentally friendly fuel resource alternatives.1,2 Lignocellulosic biomass is a sustainable source of second generation bio-fuel which can be substituted for fossil fuel.3–6 Fast pyrolysis of biomass is a highly favorable technology for production of high yield of liquid fuel (up to 70 wt%).1,2,4,7–9 Liquid fuel produced from fast pyrolysis of lignocellulosic biomass (named as pyrolysis oil or bio-oil) is considered as an economical fuel.10–12 However, bio-oil has high contents of water (15–30 wt%) and oxygen (40–55 wt%) which cause some undesirable properties such as poor chemical and thermal stability, high viscosity (poor storage stability), corrosiveness (high acidity) and low heating value.2,5,10,13–15 These negative properties of bio-oil can be removed by elimination of its oxygen content through an upgrading process. Catalytic hydrodeoxygenation (HDO) is a noteworthy method for bio-oil upgrading performed in a heterogeneous system including solid catalyst, hydrogen gas and liquid bio-oil under severe conditions (150–450 °C, 50–350 bar).1,4,6–8,11,14,16–19 In catalytic HDO of bio-oil, oxygen is eliminated as water and/or carbon oxides.1,3,4,14 Selection of a suitable catalyst is the main focus of researchers for efficient hydrotreating of bio-oil. Many studies have been conducted to develop catalysts with high catalytic activity and stability, low hydrogen consumption and high selectivity toward direct removal of oxygen.20 So far, supported and/or unsupported conventional desulfurization catalysts, noble metal and transition metal catalysts have been applied in catalytic HDO.1,2,4,10,11,14,15,21–23 Carbon,1,15,17 alumina,2,3,5,11,14 zeolites,11,13,24–26 silica,27–32 zirconia6,29,32–34 and titania6,7,34 are the common materials used as catalyst support in HDO process.

The reaction pathway for HDO of bio-oil is still unknown since it contains several hundreds of various organic compounds like acids, alcohols, aldehydes, esters, ketones, sugars, phenols and phenol derivatives resulting in hundreds of concurrent reactions during HDO of bio-oil.2,3,5,6,8,14 Therefore, investigation on the transformation of bio-oil model compounds seems to be necessary in order to attain sufficient knowledge for prediction of an overall reaction pathway for bio-oil HDO.4,14 Generally, model compounds are selected among the most reactive bio-oil compounds which lead to bio-oil instability.10 Since reactive phenolic compounds constitute the main fraction of bio-oil22,27,30,35–37 and the cleavage of Caromatic–OH bond in these compounds is difficult,37,38 many researches have been conducted in order to find suitable catalytic systems for HDO of phenols. Furthermore, phenolic compounds of bio-oil are the major cause for formation of coke precursors and catalyst deactivation during the upgrading process.8,39 Therefore, hydrodeoxygenation of bio-oil phenolic compounds to value-added chemicals is a research field of great significance.

Generally, HDO of phenolic compounds leads to formation of aromatic and alicyclic products through direct hydrogenolysis and combined hydrogenation–hydrogenolysis reaction pathways, respectively.40,41 Benzene, toluene and xylene (used as octane enhancer) are three typical aromatic compounds produced from HDO of phenols. They have several applications; apart from being used as fuel additives and solvents, they can be implemented for manufacture of various high value-added chemicals.42–44 Different parameters such as operating conditions, type of phenolic compound and catalyst properties affect reaction pathway and product selectivity in HDO of phenols.12,27 This paper gives an overview on recently used catalysts in HDO of phenol, cresol isomers, guaiacol and anisole in order to find out how catalyst type and reaction conditions could affect product selectivity.

2. HDO mechanism of phenolic compounds

2.1. HDO mechanism of phenol

Due to the very simple structure of phenol, it has extremely been applied as a model compound of bio-oil in hydrodeoxygenation process to study the reaction selectivity of different catalysts at various operating conditions. Hydroxyl functional group (–OH) is the only oxygen-containing group of phenol. Therefore, breaking the Caromatic–OH bond is the main objective of phenol deoxygenation process. Indeed, due to the high dissociation energy of Ar–OH bond (468 kJ mol−1),45 removal of the oxygen atom of hydroxyl group is very difficult which is needed to be conducted at severe operating conditions.37,46 Generally, the cleavage of Caromatic–OH bond can be occurred by dehydration or hydrodeoxygenation reactions. Dehydration is an acid-catalyzed elimination reaction which can be occurred in the case that both α- and β-carbon atoms with respect to the oxygen atom are saturated.47 Since none of these two carbon atoms are saturated in phenol molecule, dehydration reaction does not occur to eliminate the oxygen atom of phenol. In fact, catalytic hydrodeoxygenation of phenol is generally occurred in two parallel reaction pathways (Fig. 1). The first path is direct deoxygenation (DDO) or direct hydrogenolysis and contains the cleavage of Caromatic–OH bond to produce benzene as an aromatic product. The second path entails the hydrogenation of aromatic ring (HYD) which leads to formation of cyclohexanol as an intermediate, followed by a rapid dehydration and further hydrogenation to form cyclohexene and cyclohexane, respectively. Meanwhile, in catalytic hydrogenation of phenol over Pd and Pt supported on zeolites of HBeta48 and HY,49 phenol alkylation by cyclohexanol intermediate was occurred over zeolite acid sites leading to formation of bicyclic compounds (Fig. 2). Phenol and hydrogen molecules are activated by catalytic functional groups to participate in deoxygenation process. It was reported by Wang et al.21 that in HDO of phenol over Co–Mo–B catalyst, the oxygen of phenol is adsorbed on Mo4+ Brønsted acid sites and Caromatic–O bond is polarized and ruptured by nucleophilic attack of dissociative hydrogen adsorbed on the near Co active sites (Fig. 3). The free hydrogen adsorbed on the active sites of Co can contribute to deoxygenation of phenol by DDO or HYD reaction pathways. Several types of catalysts have been examined for HDO of phenol. Table 1 presents a summary of previous studies in HDO of phenol describing the effects of different catalysts and operating conditions on product selectivity.
image file: c5ra22137d-f1.tif
Fig. 1 General reaction scheme for HDO of phenol.38

image file: c5ra22137d-f2.tif
Fig. 2 Reaction pathway for conversion of phenol to bicyclic compounds over Pd/HBeta and Pt/HY catalysts.48,49

image file: c5ra22137d-f3.tif
Fig. 3 HDO mechanism of phenol over Co–Mo–O–B amorphous catalyst.46
Table 1 Overview of catalysts and operating conditions applied in HDO processing of phenol
Entry Catalyst SBET (m2 g−1) Reactor T (°C) P (bar) ta (h) Solvent Conv.b (%) Product selectivity Ref.
a Reaction time.b Conversion.c Initial pressure.d Amorphous-silica-alumina.e Phenol molar flow rate is 1.6 × 10−3 mol h−1.f Fractional conversion.g Phenol molar flow rate is 3.1 × 10−3 mol h−1.
1 Amorphous CoMoB (Mo/Co molar ratio of 3) 18.4 Batch 250 40 10 Dodecane 100 Cyclohexane (68.3 mol%) 21
Cyclohexanol (27.8)
Cyclohexene (2.2)
Benzene (0.8)
Cyclohexanone (0.7)
2 Amorphous CoMoB (Mo/Co molar ratio of 3) 18.4 Batch 275 40 10 Dodecane 100 Cyclohexane (98.2)  
Benzene (1.5)
Cyclohexene (0.3)
3 Amorphous NiMoS (Ni/(Mo + Ni) molar ratio of 0.3) 268 Batch 350 28c 1 n-Decane 96.2 Cyclohexane (52.4 mol%) 35
Benzene (30.4)
Cyclohexene (9.8)
Cyclohexanone (7.4)
4 Pd/C (5 wt% Pd) 845 Batch 200 50 0.5 Water 100 Cyclohexanol (98 C%) 27
Cyclohexanone (2)
5 Pd/C (5 wt% Pd) 845 Batch 200 50 0.5 Sodium hydroxide–water 100 Cyclohexanol (97)  
Cyclohexanone (3)
6 Pd/C (5 wt% Pd) 845 Batch 200 50 0.5 Acetic acid–water 100 Cyclohexane (75)  
Cyclohexanol (20)
Cyclohexanone (2)
7 Pd/C (5 wt% Pd) 845 Batch 200 50 0.5 Phosphoric acid–water 99 Cyclohexane (85)  
Cyclohexanol (10)
Cyclohexanone (4)
8 Pd/C (5 wt% Pd) 845 Batch 250 50 0.5 Phosphoric acid–water 100 Cyclohexane (98)  
Cyclohexanol (1)
9 Pd/C (5 wt% Pd) 845 Batch 150 50 0.5 Phosphoric acid–water 100 Cyclohexanol (74)  
Cyclohexanone (25)
10 Pd/C (5 wt% Pd) 845 Batch 150 50 1 Phosphoric acid–water 100 Cyclohexanol (94)  
Cyclohexanone (4)
11 Pd/Al2O3 (2 wt% Pd) 92 Batch 200 50 0.5 Water 82 Cyclohexanol (72)  
Cyclohexanone (28)
12 Pd/Al2O3 (2 wt% Pd) 92 Batch 200 50 0.5 Sodium hydroxide–water 72 Cyclohexanol (69)  
Cyclohexanone (31)
13 Pd/Al2O3 (2 wt% Pd) 92 Batch 200 50 0.5 Acetic acid–water 88 Cyclohexanone (56)  
Cyclohexane (20)
Cyclohexanol (20)
14 Pd/SiO2 (2 wt% Pd) 174 Batch 200 50 0.5 Water 87 Cyclohexanol (62)  
Cyclohexanone (38)
15 Pd/SiO2 (2 wt% Pd) 174 Batch 200 50 0.5 Sodium hydroxide–water 55 Cyclohexanol (53)  
Cyclohexanone (47)
16 Pd/SiO2 (2 wt% Pd) 174 Batch 200 50 0.5 Acetic acid–water 95 Cyclohexanone (64)  
Cyclohexane (20)
Cyclohexanol (13)
17 Pd/ASAd 269 Batch 200 50 0.5 Water 96 Cyclohexanol (73)  
Cyclohexanone (27)
18 Pd/ASA 269 Batch 200 50 0.5 Sodium hydroxide–water 72 Cyclohexanol (64)  
Cyclohexanone (36)
19 Pd/ASA 269 Batch 200 50 0.5 Acetic acid–water 100 Cyclohexanone (53)  
Cyclohexanol (24)
Cyclohexane (21)
20 Pt/C (5 wt% Pt) — Batch 200 50 0.5 Phosphoric acid–water 100 Cyclohexane (86)  
Cyclohexanol (13)
21 Rh/C (5 wt% Rh) — Batch 200 50 0.5 Phosphoric acid–water 100 Cyclohexane (92)  
Cyclohexanol (7)
22 Ru/C (5 wt% Ru) — Batch 200 50 0.5 Phosphoric acid–water 100 Cyclohexane (88)  
Cyclohexanol (10)
Cyclohexanone (1)
23 Ru/C (5 wt% Ru) 717 Batch 250 100 4.3 — — Cyclohexane (∼45 wt%) 8
Cyclohexanol (∼43)
24 Ru/C prepared from Ru(acac)3 1% — Batch 250 100 4.3 — — Cyclohexane (∼4)  
Cyclohexanol (∼72)
Phenol (∼12)
25 Ru/C prepared from Ru(acac)3 3% — Batch 250 100 4.3 — — Cyclohexane (∼2)  
Cyclohexanol (∼98)
26 Ru/C prepared from Ru(acac)3 5% 774 Batch 250 100 4.3 — — Cyclohexane (∼42)  
Cyclohexanol (∼47)
27 Ru/C prepared from RuCl3 1%   Batch 250 100 4.3 — — Cyclohexane (∼9)  
Cyclohexanol (∼81)
28 Ru/C prepared from RuCl3 3% — Batch 250 100 4.3 — — Cyclohexane (∼42)  
Cyclohexanol (∼58)
29 Ru/C prepared from Ru(NO3)3 3% — Batch 250 100 4.3 — — Cyclohexane (∼12)  
Cyclohexanol (∼82)
30 Ru/C prepared from Ru(NO3)3 5% — Batch 250 100 4.3 — — Cyclohexane (∼15)  
Cyclohexanol (∼75)
31 CoMoS/Al2O3 235 Continuous 300 28.5 — n-Heptane or n-decane 30 Aromatic/hydrogenated compounds ratio (∼0.2) 40
32 CoMoS/Al2O3 — Batch 300 50 4 Dodecane 27 Benzene (37 mol% of product) 50
Cyclohexene, cyclohexanol, methylcyclohexane, and methylcyclohexene (4)
33 Amorphous NiMoB — Batch 250 40 10 Decane 98.5 Cyclohexanol (72.5%) 38
Cyclohexanone (0.7)
Benzene (5)
Cyclohexane (21.3)
Cyclohexene (0.4)
34 Amorphous CoNiMoB (Co molar ratio of 150) — Batch 250 40 10 Decane 98 Cyclohexanol (6.5)  
Cyclohexanone (0.3)
Benzene (3.2)
Cyclohexane (89.8)
Cyclohexene (0.1)
35 Amorphous CoNiMoB (Co molar ratio of 225) — Batch 250 40 10 Decane 98.5 Cyclohexanol (19.3)  
Cyclohexanone (1.2)
Benzene (0.5)
Cyclohexane (76.8)
Cyclohexene (2.2)
36 Amorphous CoNiMoB (Co molar ratio of 300) — Batch 250 40 10 Decane 62.3 Cyclohexanol (12.5)  
Cyclohexanone (1.4)
Benzene (0.5)
Cyclohexane (40.4)
Cyclohexene (45.2)
37 CoMoS/Al2O3 — Continuous 250 15 24 — 36.4 Benzene (92.9%) 41
Cyclohexane and cyclohexene (7.1)
38 CoMoS/Al2O3 — Continuous 300 15 7 — 71.9 Benzene (86.4)  
Cyclohexane and cyclohexene (13.6)
39 CoMoO/Al2O3 — Continuous 250 15 6 — 0.12 Benzene (100)  
40 Amorphous NiWB 40.2 Batch 225 40 4 Dodecane 100 Cyclohexanol (76.4 mol% of product) 37
Cyclohexanone (1.2)
Cyclohexane (22.3)
Cyclohexene (0.1)
41 Amorphous LaNiWB (La molar ratio of 50) 54.7 Batch 225 40 4 Dodecane 97.3 Cyclohexanol (40.6)  
Cyclohexanone (0.8)
Cyclohexane (37.2)
Cyclohexene (21.3)
42 Amorphous LaNiWB (La molar ratio of 150) 25.5 Batch 225 40 4 Dodecane 69 Cyclohexanol (11.8)  
Cyclohexanone (9.1)
Cyclohexane (22.9)
Cyclohexene (56.2)
43 CoMoS/MgO 45.5 Batch 350 50c 1 Supercritical hexane 16.74 Benzene (0.01 GC area%) 16
Cyclohexanol (1.12)
Cyclohexanone (0.75)
44 CoMoPS/MgO 51.1 Batch 350 50c 1 Supercritical hexane 35.16 Benzene (13.23)  
45 RANEY® Ni with Nafion/SiO2 140 Batch 300 40 2 Water 100 Cyclohexane (93 C%) 51
Cyclohexene (3.9)
Benzene (2.9)
46 Amorphous CoMoOB (2.5 h preparation time) 22.3 Batch 275 40 10 Dodecane 97.9 Cyclohexanol (∼5 mol%) 46
Cyclohexanone (∼2)
Cyclohexane (∼80)
Benzene (∼3)
Cyclohexene (∼10)
47 Amorphous CoMoOB (4 h preparation time) 20.5 Batch 275 40 10 Dodecane 99.5 Cyclohexanol, cyclohexanone and cyclohexene (∼3)  
Cyclohexane (∼95)
Benzene (∼2)
48 Amorphous CoMoOB (5 h preparation time) 17.2 Batch 275 40 10 Dodecane 100 Cyclohexanol, cyclohexanone and cyclohexene (∼2)  
Cyclohexane (∼96)
Benzene (∼2)
49 Oxide NiWP/AC 1099 Continuous 150 15 — n-Octane 56 Cyclohexane (∼78%) 52
Cyclohexene (∼3)
Cyclohexanol (∼19)
50 Oxide NiWP/AC 1099 Continuous 300 15 — n-Octane 98 Cyclohexane (∼95)  
Benzene (∼4)
Methylcyclopentane (∼1)
51 Oxide NiWSi/AC 1071 Continuous 150 15 — n-Octane 56 Cyclohexane (∼86)  
Cyclohexene (∼3)
Cyclohexanol (∼11)
52 Oxide NiWSi/AC 1071 Continuous 300 15 — n-Octane 98 Cyclohexane (∼90)  
Benzene (∼6.5)
Methylcyclopentane (∼3.5)
53 Amorphous NiMoB (Mo/Ni molar ratio of 1) 143.5 Batch 225 40 7 Dodecane 100 Cyclohexanol (74 mol%) 53
Cyclohexanone (0.1)
Benzene (5.3)
Cyclohexane (20.6)
54 Amorphous NiMoB (Mo/Ni molar ratio of 2) 122.4 Batch 225 40 7 Dodecane 72.1 Cyclohexanol (19.2)  
Cyclohexanone (0.7)
Benzene (8.9)
Cyclohexene (0.1)
Cyclohexane (71)
55 Amorphous NiMoB (Mo/Ni molar ratio of 3) 38.4 Batch 225 40 7 Dodecane 94.9 Cyclohexanol (86.2)  
Cyclohexanone (0.8)
Benzene (3.8)
Cyclohexene (0.1)
Cyclohexane (9.1)
56 Pd/C 845 Batch 80 50 7 Phosphoric acid–water ∼45 Cyclohexanol (∼98 C%) 54
Cyclohexanone (∼2)
57 Pd/C with HZSM-5 1062 Batch 200 50 2 Water 100 Cycloalkanes (100 C%) 55
58 Crystalline MoS 11 Batch 350 28c 1 n-Decane 30 Cyclohexanone (16.2 mol%) 56
Benzene (20.3)
Cyclohexane (34.5)
Cyclohexene (29)
59 Amorphous MoS 368 Batch 350 28c 1 n-Decane 71 Cyclohexanone (4.5)  
Benzene (65.6)
Cyclohexane (22)
Cyclohexene (7.9)
60 Amorphous CoMoS (Co/(Mo + Co) molar ratio of 0.2) 275 Batch 350 28c 1 n-Decane 98.2 Cyclohexanone (3.2)  
Benzene (80.3)
Cyclohexane (12)
Cyclohexene (4.5)
61 Pd/CeO2 (1 wt% Pd) 229 Continuous 180 1 1 Benzene 50.2 Cyclohexanone (72.3%) 57
Cyclohexanol (25.4)
Cyclohexane (2.3)
62 Pd/CeO2 (1 wt% Pd) 229 Continuous 180 1 1 Toluene 52.8 Cyclohexanone (73.3)  
Cyclohexanol (23.9)
Cyclohexene (2.8)
63 Pd/CeO2 (3 wt% Pd) 166 Continuous 180 1 1 Benzene 81.4 Cyclohexanone (46.2)  
Cyclohexanol (34.8)
Cyclohexane (19)
64 Pd/CeO2 (3 wt% Pd) 166 Continuous 180 1 1 Toluene 81.8 Cyclohexanone (46.3)  
Cyclohexanol (48.9)
Cyclohexene (4.8)
65 Pd/CeO2 (3 wt% Pd) 166 Continuous 180 1 1 Cyclohexane 93.7 Cyclohexanone (47)  
Cyclohexanol (53)
66 Pd/CeO2 (3 wt% Pd) 166 Continuous 180 1 1 Ethanol 28 Cyclohexanone (86.8)  
Cyclohexanol (13.2)
67 Pd/ZrO2 (1 wt% Pd) 269 Continuous 180 1 1 Benzene 52.7 Cyclohexanone (92.2)  
Cyclohexanol (5.1)
68 Pd/ZrO2 (3 wt% Pd) 254 Continuous 180 1 1 Benzene 62.9 Cyclohexanone (93)  
Cyclohexanol (3.5)
Cyclohexane (2.7)
69 Pd/MgO (1 wt% Pd) — Continuous 180 1 1 Benzene 77 Cyclohexanone (90.5)  
Cyclohexanol (4.9)
Cyclohexane (0.6)
70 Pd/Al2O3 (1 wt% Pd) — Continuous 180 1 1 Benzene 10 Cyclohexanone (100)  
71 Pd/TiO2 (3 wt% Pd) — Continuous 180 1 1 Benzene 16.7 Cyclohexanone (98.8)  
Cyclohexanol (1.2)
72 5% w/w Pd/SiO2 prepared from Pd(NO3)2e Surface area of SiO2 > 200 Continuous 150 1 — — 0.27f Cyclohexanone/cyclohexanol ratio (1.9) 58
73 5% w/w Pd/SiO2 prepared from Pd(NO3)2g Surface area of SiO2 > 200 Continuous 150 1 — — 0.13f Cyclohexanone/cyclohexanol ratio (2.1)  
74 5% w/w Pd/SiO2 prepared from Pd(C2H3O2)2e Surface area of SiO2 > 200 Continuous 150 1   — 0.17f Cyclohexanone/cyclohexanol ratio (10.6)  
75 5% w/w Pd/SiO2 prepared from Pd(C2H3O2)2g Surface area of SiO2 > 200 Continuous 150 1 — — 0.06f Cyclohexanone/cyclohexanol ratio (46.6)  
76 Pd–Yb/SiO2e Surface area of SiO2 > 200 Continuous 150 1 — — 0.86f Cyclohexanone/cyclohexanol ratio (0.7)  
77 Pd–Yb/SiO2g Surface area of SiO2 > 200 Continuous 150 1 — — 0.65f Cyclohexanone/cyclohexanol ratio (1.3)  

2.2. HDO mechanism of cresol

Cresol is another phenolic compound which has been used as model compound for study of bio-oil HDO. Based on the position of methyl group, o-cresol, m-cresol and p-cresol are the three isomers of cresol molecule. Like phenol, hydroxyl group (–OH) is the only oxygen-containing group of cresol. Hydrogenolysis of Caromatic–OH bond producing toluene (as aromatic product) and ring hydrogenation of cresol followed by rapid dehydration forming methylcyclohexane (as methylcycloalkane product) are the two main reaction pathways occurred in HDO of cresol (Fig. 4).59 Toluene and methylcyclohexane which have high octane number are the two valuable products formed from HDO of cresol.36 In comparison with phenol molecule, the presence of methyl group in cresol molecule favors aromatic production in HDO process.40 Various isomers of cresol behave differently in hydrodeoxygenation reaction. Massoth et al.40,41 reported that o-cresol had lower conversion compared to m- and p-cresol probably due to the steric hindrance to adsorption of o-cresol with methyl group adjacent to the hydroxyl group. Meanwhile, type of catalyst could change overall pathway of cresols HDO. The lowest unoccupied molecular orbital of Caromatic–OH is antibonding and the surfaces which transfer electron to this orbital promote the cleavage of Caromatic–OH bond.22 Catalysts with high electron density facilitate the above mentioned electron transfer improving catalytic activity. An overview of catalysts and operating conditions applied in HDO of cresols is presented in Table 2.
image file: c5ra22137d-f4.tif
Fig. 4 Simple reaction scheme of HDO of p-cresol.59
Table 2 Overview of catalysts and operating conditions applied in HDO processing of cresols
Entry Catalyst SBET (m2 g−1) Cresol isomers Reactor T (°C) P (bar) ta (h) Solvent Conv.b (%) Product selectivity Ref.
a Reaction time.b Conversion.
1 Pt/C 764 p-Cresol Batch 300 83 1 Water ∼98 Toluene (∼45 C%) 36
4-Methylcyclohexanol (∼2)
Methylcyclohexane (∼53)
2 Pt/C 764 p-Cresol Batch 300 83 1 n-Heptane ∼99 Toluene (∼2)  
4-Methylcyclohexanol (∼92)
Methylcyclohexane (∼6)
3 Pt/Al2O3 106 p-Cresol Batch 300 83 1 Water ∼96 Toluene (∼50)  
4-Methylcyclohexanol (∼22)
Methylcyclohexane (∼28)
4 Pt/Al2O3 106 p-Cresol Batch 300 83 1 n-Heptane ∼100 4-Methylcyclohexanol (∼20)  
Methylcyclohexane (∼80)
5 Pd/C — p-Cresol Batch 300 83 1 Water ∼69 Toluene (∼20)  
4-Methylcyclohexanol (∼64)
Methylcyclohexane (∼16)
6 Pd/C — p-Cresol Batch 300 83 1 n-Heptane ∼99 4-Methylcyclohexanol (∼88)  
Methylcyclohexane (∼12)
7 Pd/Al2O3 — p-Cresol Batch 300 83 1 Water ∼30 Toluene (∼9)  
4-Methylcyclohexanol (∼90)
Methylcyclohexane (∼1)
8 Pd/Al2O3 — p-Cresol Batch 300 83 1 n-Heptane ∼99 4-Methylcyclohexanol (∼30)  
Methylcyclohexane (∼70)
9 Ru/C — p-Cresol Batch 300 83 1 Water ∼50 Toluene (∼12)  
4-Methylcyclohexanol (∼68)
Methylcyclohexane (∼6)
10 Ru/C — p-Cresol Batch 300 83 1 n-Heptane ∼100 4-Methylcyclohexanol (∼78)  
Methylcyclohexane (∼18)
11 Ru/Al2O3 — p-Cresol Batch 300 83 1 Water ∼30 Toluene (∼12)  
4-Methylcyclohexanol (∼72)
Methylcyclohexane (∼4)
12 Ru/Al2O3 — p-Cresol Batch 300 83 1 n-Heptane ∼94 Toluene (∼7)  
4-Methylcyclohexanol (∼54)
Methylcyclohexane (∼22)
13 CoMo/Al2O3 — p-Cresol Batch 300 83 1 n-Heptane ∼35 Toluene (∼98)  
Methylcyclohexane (∼2)
14 NiMo/Al2O3 — p-Cresol Batch 300 83 1 n-Heptane ∼20 Toluene (∼26)  
Methylcyclohexane (∼74)
15 NiW/Al2O3 — p-Cresol Batch 300 83 1 n-Heptane ∼10 Toluene (∼46)  
Methylcyclohexane (∼54)
16 MoP 8.8 p-Cresol Batch 325 41.4 5 Decalin ∼60 Toluene (∼58%) 22
Hydrogenated products (∼42)
17 MoP 8.8 p-Cresol Batch 350 44 5 Decalin ∼94 Toluene (∼50)  
Hydrogenated products (∼50)
18 MoS2 4.3 p-Cresol Batch 325 41.4 5 Decalin ∼35 Toluene (∼70)  
Hydrogenated products (∼30)
19 MoS2 4.3 p-Cresol Batch 350 44 5 Decalin ∼60 Toluene (∼82)  
Hydrogenated products (∼15)
Isomerization products (∼3)
20 MoO2 4.8 p-Cresol Batch 325 41.4 5 Decalin ∼34 Toluene (∼64)  
Hydrogenated products (∼36)
21 MoO2 4.8 p-Cresol Batch 350 44 5 Decalin ∼60 Toluene (∼68)  
Hydrogenated products (∼26)
Isomerization products (∼6)
22 MoO3 0.3 p-Cresol Batch 325 41.4 5 Decalin ∼80 Toluene (∼60)  
Hydrogenated products (∼32)
Isomerization products (∼8)
23 MoO3 0.3 p-Cresol Batch 350 44 5 Decalin ∼100 Toluene (∼60)  
Hydrogenated products (∼18)
Isomerization products (∼22)
24 MoP 8 p-Cresol Batch 350 44 5 Decalin 45 Toluene (49%) 60
Methylcyclohexane (50)
Dimethylcyclopentane (1)
25 MoP prepared with citric acid and calcinated at 550 °C 112 p-Cresol Batch 350 44 5 Decalin 71 Toluene (51)  
Methylcyclohexane (47)
Dimethylcyclopentane (2)
26 MoP prepared with citric acid and calcinated at 600 °C 75 p-Cresol Batch 350 44 5 Decalin 56 Toluene (56)  
Methylcyclohexane (42)
Dimethylcyclopentane (2)
27 Pt/Al2O3 (0.5 wt% Pt) — m-Cresol Continuous 260 1 — — 9.3 Methylcyclohexanol (19.4%) 61
Toluene (61.6)
Methylcyclohexane (17.6)
Phenol (1.4)
28 Pt/Al2O3 (1.0 wt% Pt) — m-Cresol Continuous 260 1 — — 18 Methylcyclohexanol (16.4)  
Toluene (69.7)
Methylcyclohexane (12.9)
Phenol (0.6)
29 Pt/Al2O3 (1.7 wt% Pt) — m-Cresol Continuous 260 1 — — 38.3 Methylcyclohexanol (16.6)  
Toluene (66.8)
Methylcyclohexane (13.7)
Methylcyclohexene (0.3)
Phenol (0.2)
30 Pt/F–Al2O3 — m-Cresol Continuous 260 1 — — 63 Methylcyclohexanol (1.4)  
Toluene (82.3)
Methylcyclohexane (16)
Phenol (0.3)
31 Pt/K–Al2O3 — m-Cresol Continuous 260 1 — — 25.7 Methylcyclohexanol (16.3)  
Toluene (77.6)
Methylcyclohexane (4.4)
Methylcyclohexene (1.1)
Phenol (0.5)
32 K–Pt/Al2O3 — m-Cresol Continuous 260 1 — — 25.1 Methylcyclohexanol (11.7)  
Toluene (73)
Methylcyclohexane (14.4)
Methylcyclohexene (0.8)
33 Pt/SiO2 — m-Cresol Continuous 260 1 — — 54.9 Methylcyclohexanol (16.6)  
Toluene (77.9)
Methylcyclohexane (4.2)
Methylcyclohexene (1)
Phenol (0.2)
34 Pt/K–SiO2 — m-Cresol Continuous 260 1 — — 13.5 Methylcyclohexanol (84.6)  
Toluene (5)
Methylcyclohexane (6.3)
Methylcyclohexene (3.4)
Phenol (0.6)
35 K–Pt/SiO2 — m-Cresol Continuous 260 1 — — 16.4 Methylcyclohexanol (90)  
Toluene (1.3)
Methylcyclohexane (5)
Methylcyclohexene (3.1)
Phenol (0.5)
36 Ga/HBeta (2.9 wt% Ga) 580 m-Cresol Continuous 450 1 — — 82.78 Benzene (12.62%) 62
Toluene (27.85)
Xylene (11.88)
Other hydrocarbons C2–C6 (2.22)
Phenol (16.65)
Oxygen-containing compounds (26.06)
37 Ga/SiO2 (2.9 wt% Ga) 370 m-Cresol Continuous 450 1 — — 4.49 Benzene (3.40)  
Toluene (7.19)
Phenol (62.88)
Oxygen-containing compounds (9.08)
38 Ga/HZSM-5 (2.8 wt% Ga) 510 m-Cresol Continuous 450 1 — — 40.32 Benzene (3.08)  
Toluene (22.06)
Xylene (2.13)
Other hydrocarbons C2–C6 (2.23)
Phenol (49.38)
Oxygen-containing compounds (20.29)
39 CoMoS/Al2O3 235 m-Cresol Continuous 300 28.5 — — ∼25 Aromatic/hydrogenated compounds ratio (∼0.6) 40
40 CoMoS/Al2O3 235 p-Cresol Continuous 300 28.5 — — ∼27 Aromatic/hydrogenated compounds ratio (∼0.75)  
41 CoMoS/Al2O3 — o-Cresol Batch 300 50 4 Dodecane 23 Toluene (65 mol% of product) 50
Cyclohexene, cyclohexanol, methylcyclohexane, and methylcyclohexene (3)
42 CoMoS/Al2O3 — p-Cresol Batch 300 50 4 Dodecane 22 Toluene (60)  
Cyclohexene, cyclohexanol, methylcyclohexane, and methylcyclohexene (6)

2.3. HDO mechanism of guaiacol

Guaiacol with two oxygen functionalities (phenolic functional group (–OH) and methoxy functional group (–OCH3)) is an attractive model compound of bio-oil for HDO survey.63–65 The Caromatic–OH bond of phenolic group is stronger than the C–O bond of methoxy group, and its rupture is happened at highly strict conditions (high temperature and pressure). In HDO of guaiacol, methoxy group undergoes two reactions (Fig. 5): demethylation (DME) which is occurred by the cleavage of C–O bond of –OCH3 group leading to production of catechol and methane (as byproduct), and direct demethoxylation (DMO) which is happened by the rupture of Caromatic–O bond resulting in formation of phenol and methanol (as byproduct). Oxygen atom of phenolic group is also eliminated through two different routes: cleavage of the bond between aromatic ring carbon atom and phenol group oxygen atom as well as aromatic ring hydrogenation followed by deletion of –OH group.29,32 Guaiacol type molecules in bio-oil are the main agents for coke formation.66 Hydrotreating of guaiacol without presence of catalyst yields oxygen-containing compounds such as catechol, phenol and cresol indicating the necessity of catalyst use for complete deoxygenation of guaiacol.67–69 Product selectivities obtained from HDO of guaiacol at different reaction conditions is illustrated in Table 3.
image file: c5ra22137d-f5.tif
Fig. 5 General scheme of guaiacol conversion pathways.32
Table 3 Overview of catalysts and operating conditions applied in HDO processing of guaiacol
Entry Catalyst SBET (m2 g−1) Reactor T (°C) P (bar) ta (h) Conv.b (%) Product selectivity Ref.
a Reaction time.b Conversion.c Initial pressure.
1 Ni/SiO2 (64.2 wt% Ni) 38 Batch 320 170 1 96.7 Saturated cyclic C5–C7 hydrocarbons and benzene (3.6 mol%) 29
Saturated cyclic oxygen-containing C5 and C6 hydrocarbons (95.9)
Phenol and anisole (0.5)
2 Ni/SiO2 (55.4 wt% Ni) 216 Batch 320 170 1 97.5 Saturated cyclic C5–C7 hydrocarbons and benzene (93.1)  
Saturated cyclic oxygen-containing C5 and C6 hydrocarbons (1.5)
Phenol and anisole (2.2)
3 NiCuLa/ZrO2–SiO2 66 Batch 320 170 1 85.6 Saturated cyclic C5–C7 hydrocarbons and benzene (92.9)  
Saturated cyclic oxygen-containing C5 and C6 hydrocarbons (3)
Phenol and anisole (2.2)
4 NiCu/CeO2–ZrO2 82 Batch 320 170 1 94.2 Saturated cyclic C5–C7 hydrocarbons and benzene (1.7)  
Saturated cyclic oxygen-containing C5 and C6 hydrocarbons (97.9)
Phenol and anisole (0.4)
5 NiCu/Al2O3 109 Batch 320 170 1 80.2 Saturated cyclic C5–C7 hydrocarbons and benzene (61.8)  
Saturated cyclic oxygen-containing C5 and C6 hydrocarbons (35.6)
Phenol and anisole (1.2)
6 NiCu/SiO2 142 Batch 320 170 1 87 Saturated cyclic C5–C7 hydrocarbons and benzene (83.6)  
Saturated cyclic oxygen-containing C5 and C6 hydrocarbons (8.2)
Phenol and anisole (3.5)
7 Ni2P/SiO2 309 Continuous 300 1 — 80 Phenol (30%) 31
Benzene (60)
Anisole (10)
8 Co2P/SiO2 307 Continuous 300 1 — 70 Phenol (32)  
Benzene (52)
Anisole (1)
C3–C5 hydrocarbons (15)
9 Fe2P/SiO2 233 Continuous 300 1 — 64 Phenol (94)  
Anisole (6)
10 WP/SiO2 147 Continuous 300 1 — 60 Phenol (100)  
11 MoP/SiO2 207 Continuous 300 1 — 54 Phenol (28)  
Benzene (53)
Toluene (4)
C3–C5 hydrocarbons (15)
12 Pt/ZrO2 17 Batch 100 80 5 10 Cyclohexanol (∼65 mol%) 33
1-Methyl-1,2-cyclohexanediol (∼15)
1,2-Dimethoxybenzene (∼15)
Cyclohexane (∼1)
13 Pd/ZrO2 17 Batch 100 80 5 13.7 Cyclohexanol (∼40%)  
1-Methyl-1,2-cyclohexanediol (∼37)
1,2-Dimethoxybenzene (∼20)
Cyclohexane (∼2)
14 Rh/ZrO2 20 Batch 100 80 5 98.9 Cyclohexanol (∼12)  
1-Methyl-1,2-cyclohexanediol (∼72)
1,2-Dimethoxybenzene (∼3)
15 PdPt/ZrO2 16 Batch 100 80 5 5.2 Cyclohexanol (∼73)  
1-Methyl-1,2-cyclohexanediol (∼4)
1,2-Dimethoxybenzene (∼17)
16 RhPd/ZrO2 21 Batch 100 80 5 32.7 Cyclohexanol (∼32)  
1-Methyl-1,2-cyclohexanediol (∼44)
1,2-Dimethoxybenzene (∼8)
Cyclohexane (∼0.05)
17 RhPt/ZrO2 23 Batch 100 80 5 98.7 Cyclohexanol (∼7)  
1-Methyl-1,2-cyclohexanediol (∼91)
18 CoMoS/Al2O3 — Batch 100 80 5 13.8 Cyclohexanol (∼73)  
1,2-Dimethoxybenzene (∼17)
19 MoS2 82 Continuous 300 40 — ∼55 Cyclohexane (∼25%) 23
Methylcyclopentane (∼16)
Benzene (∼18)
Cyclohexene (trace)
Cyclohexylbenzene (trace)
20 MoS2/Al2O3 250 Continuous 300 40 — ∼13 Cyclohexene (∼5.8)  
Benzene (∼1.8)
Cyclohexylbenzene (∼1)
Methyl-containing products (∼4.8)
Cyclohexane (∼1)
21 CoMoS 110 Continuous 300 40 — ∼50 Benzene (∼42)  
Cyclohexene (∼2)
Cyclohexane (∼6)
Toluene (∼2)
22 CoMoS/Al2O3 230 Continuous 300 40 — ∼13 Methyl-containing products (∼6.3)  
Cyclohexene (∼2.8)
Benzene (∼1.8)
Cyclohexane (∼2)
23 MoS2/Al2O3 230 Continuous 300 40 — ∼25 Catechol (∼51%) 34
Methylcatechol (∼18)
Phenol (∼9)
Cresol (∼9)
Oxygen-free products (∼2)
Heavy compounds (∼7)
Light compounds (∼4)
24 MoS2/TiO2 120 Continuous 300 40 — ∼25 Catechol (∼20)  
Phenol (∼60)
Cresol (∼7)
Oxygen-free products (∼4)
Light compounds (∼9)
25 MoS2/ZrO2 94 Continuous 300 40 — ∼25 Catechol (∼37)  
Phenol (∼45)
Cresol (∼4)
Oxygen-free products (∼5)
Light compounds (∼9)
26 CoMoS/Al2O3 190 Continuous 300 40 — ∼90 Catechol (∼10)  
Methylcatechol (∼4)
Phenol (∼34)
Cresol (∼28)
Oxygen-free products (∼15)
Heavy compounds (∼4)
Light compounds (∼5)
27 CoMoS/TiO2 112 Continuous 300 40 — 100 Catechol (∼1)  
Methylcatechol (trace)
Phenol (∼61)
Cresol (∼6)
Oxygen-free products (∼18)
Light compounds (∼14)
28 CoMoS/ZrO2 90 Continuous 300 40 — 100 Catechol (∼1)  
Phenol (∼55)
Cresol (∼2)
Oxygen-free products (∼30)
Light compounds (∼12)
29 Ni/SiO2 216 Batch 320 170 1 97.5 Aliphatics C5–C7 and benzene (90.9 mol%) 32
Oxygen-containing aliphatics C5–C7 (1.4)
Phenol, anisole and methoxy-methylphenol (2.1)
Products of aromatic ring condensation (3.1)
30 NiCu/SiO2 142 Batch 320 170 1 87.1 Aliphatics C5–C7 and benzene (72.7)  
Oxygen-containing aliphatics C5–C7 (7.1)
Phenol, anisole and methoxy-methylphenol (3)
Products of aromatic ring condensation (4.2)
31 NiCu/CeO2–ZrO2 82 Batch 320 170 1 94.2 Aliphatics C5–C7 and benzene (1.6)  
Oxygen-containing aliphatics C5–C7 (92.5)
Phenol, anisole and methoxy-methylphenol (0.4)
32 NiCu/Al2O3 109 Batch 320 170 1 80.3 Aliphatics C5–C7 and benzene (49.9)  
Oxygen-containing aliphatics C5–C7 (28.9)
Phenol, anisole and methoxy-methylphenol (0.9)
Products of aromatic ring condensation (1.1)
33 NiCu/ZrO2–SiO2–La2O3 66 Batch 320 170 1 85.6 Aliphatics C5–C7 and benzene (80.1)  
Oxygen-containing aliphatics C5–C7 (2.7)
Phenol, anisole and methoxy-methylphenol (1.9)
Products of aromatic ring condensation (1.6)
34 Pt/Al2O3 120 Batch 250 40c 1 100 Yield of product (mol%) 70
Cyclohexane (23)
Cyclohexanol (19)
Cyclohexanone (6)
2-Methoxycyclohexanol (26)
35 Pt/SiO2–Al2O3 626 Batch 250 40c 1 50 Cyclohexane (17)  
36 Pt/nitric-acid-treated carbon black 290 Batch 250 40c 1 100 Cyclohexane (14)  
Cyclohexanol (15)
Cyclohexanone (1)
2-Methoxycyclohexanol (44)
37 Ru/Al2O3 119 Batch 250 40c 1 100 Cyclohexane (22)  
Cyclohexanol (35)
2-Methoxycyclohexanol (7)
38 Ru/SiO2–Al2O3 518 Batch 250 40c 1 100 Cyclohexane (60)  
Cyclohexanol (8)
39 Ru/nitric-acid-treated carbon black 210 Batch 250 40c 1 100 Cyclohexane (12)  
Cyclohexanol (19)
2-Methoxycyclohexanol (60)
40 Rh/Al2O3 122 Batch 250 40c 1 100 Cyclohexane (25)  
Cyclohexanol (8)
Cyclohexanone (1)
2-Methoxycyclohexanol (9)
41 Rh/SiO2–Al2O3 569 Batch 250 40c 1 100 Cyclohexane (57)  
42 Rh/nitric-acid-treated carbon black 238 Batch 250 40c 1 100 Cyclohexane (20)  
Cyclohexanol (6)
Cyclohexanone (1)
2-Methoxycyclohexanol (41)
43 Pd/Al2O3 114 Batch 250 40c 1 100 Cyclohexane (19)  
Cyclohexanol (13)
2-Methoxycyclohexanol (13)
44 Pd/SiO2–Al2O3 520 Batch 250 40c 1 100 Cyclohexane (46)  
45 Pd/nitric-acid-treated carbon black 240 Batch 250 40c 1 54 Cyclohexane (1)  
Cyclohexanol (1)
2-Methoxycyclohexanol (41)
46 MoN/Al2O3 (ammonia as nitriding agent) 191 Batch 300 50 4 66 At 10% guaiacol conversion: phenol (1%), catechol (26) 20
47 MoN/Al2O3 (N2/H2 as nitriding agent) 183 Batch 300 50 4 62 At 10% guaiacol conversion: phenol (2), catechol (26)  
48 MoN/SBA-15 (ammonia as nitriding agent) 418 Batch 300 50 4 44 At 10% guaiacol conversion: phenol (26), catechol (6)  
49 MoN/SBA-15 (N2/H2 as nitriding agent) 397 Batch 300 50 4 40 At 10% guaiacol conversion: phenol (22), catechol (9)  
50 CoMoN/SBA-15 (ammonia as nitriding agent) 387 Batch 300 50 4 24 At 10% guaiacol conversion: phenol (12%), catechol (6%)  
51 CoMoN/Al2O3 (ammonia as nitriding agent) 182 Batch 300 50 4 58 At 10% guaiacol conversion: phenol (5), catechol (22)  
52 Re/ZrO2 (sulfided in H2/H2S mixture) 77 Batch 300 50 ∼4 ∼50 Yield of product (%): catechol (∼10), phenol (∼34), cyclohexane (∼4), cyclohexene (∼4), benzene (∼1) 74
53 Re/ZrO2 (sulfided in N2/H2S mixture) 77 Batch 300 50 ∼4 ∼70 Catechol (∼18)  
Phenol (∼46)
Cyclohexane (∼4)
Cyclohexene (∼5)
Benzene (∼2)
54 Re/ZrO2 (sulfated support) 79 Batch 300 50 ∼4 ∼10 Catechol (∼7)  
Phenol (∼3)
Cyclohexane (∼1)
55 Ni2P/Al2O3 90 Continuous 300 1 — 99.6 Yield of product (C%): benzene (30.9), phenol (13.7), cresol (1.5) 75
56 Ni2P/ZrO2 47 Continuous 300 1 — 96.5 Benzene (32.4)  
Phenol (25.5)
Cresol (1.3)
Catechol (0.3)
57 Ni2P/SiO2 127 Continuous 300 1 — 99.5 Benzene (71.9)  
Phenol (1.9)
Cresol (0.1)
58 MoS2/C 126 Batch 300 50c 5 ∼55 Phenol (52 mol%) 69
Cyclohexane (4.2)
Cyclohexene (8)
Cyclohexanol (5)
Anisole (0.3)
Cresol (1.2)
Benzene (0.4)
Catechol (1.8)
59 Pt/MgO — Continuous 250 69 — 62.9 Cyclopentane (25%) 76
Cyclohexanol (39)
Cyclohexanone (24)
Phenol (0.6)
Methylcyclohexanol (0.6)
1,2-Dimethoxybenzene (1)

To study the role of catalyst type in reaction selectivity of guaiacol HDO, hydrodeoxygenation of guaiacol over Pt, Rh, Pd and Ru noble metal catalysts supported on Al2O3, SiO2–Al2O3 and nitric acid treated carbon black (NAC) were conducted in a batch reactor at 250 °C and 40 bar hydrogen pressure.70 Cyclohexane selectivity in HDO of guaiacol over these catalysts demonstrates a different reaction route from the common pathway of this process (guaiacol → phenol → hydrodeoxygenated products).20,31,42,71–73 Phenol was not produced and 2-methoxycyclohexanol, cyclohexanol and cyclohexanone were intermediate products (Fig. 6). Non-production of phenol was attributed to the high hydrogenation capability of noble metal catalysts. In fact, the mechanism of guaiacol HDO in this study included two sequential steps; hydrogenation of guaiacol benzene ring followed by demethoxylation and dehydroxylation of oxygenates. Noble metals hydrogenate the aromatic ring and acid sites of support catalyze hydrogenolysis reaction. However, at different experimental conditions (continuous reactor at 300 °C, 1.4 bar) over Pt/Al2O3, it was shown that the cleavage of Caromatic–O bond of guaiacol due to direct hydrogenolysis led to phenol and anisole as primary products.26 The high selectivity toward phenol shows that demethoxylation was more favorable than dehydroxylation in direct deoxygenation pathway. Meanwhile, presence of methylated compounds such as 3-methylcatechol and methylguaiacols indicates that transalkylation (methyl group transfer) route was also occurred in HDO of guaiacol over Pt/Al2O3. Methylation reaction is taken place on the acidic alumina support.


image file: c5ra22137d-f6.tif
Fig. 6 Proposed reaction pathway for conversion of guaiacol to cyclohexane over Pt, Rh, Pd, and Ru catalysts supported on Al2O3, SiO2–Al2O3, and NAC.70

2.4. HDO mechanism of anisole

Anisole with an isolated methoxy group (–OCH3) is another phenolic compound applied for study of bio-oil HDO. A bifunctional catalyst consisting acid and metal is required for hydrogenolysis and hydrocracking of Caromatic–O–Cmethyl bonds in anisole molecule.77 HDO of anisole could proceed via two common reaction pathways (Fig. 7): (i) demethylation of anisole to phenol followed by hydrogenolysis and ring hydrogenation of phenol to benzene and cyclohexane, (ii) transalkylation to cresols, toluene and xylenols. Since Caromatic–O bond is stronger than Cmethyl–O bond in anisole molecule, the possibility of direct cleavage of methoxy group is weak.78 In a study held for anisole HDO over sulfided CoMo/Al2O3 catalyst, it was inferred that the two reactions of demethylation and methyl transfer are occurred on different active sites of catalyst.40,41 Generally, acidic sites of catalyst participate in anisole demethylation. An overview of products selectivity obtained by HDO of anisole using different types of catalyst and operating conditions is presented in Table 4.
image file: c5ra22137d-f7.tif
Fig. 7 HDO network of anisole.41
Table 4 Overview of catalysts and operating conditions applied in HDO processing of anisole
Entry Catalyst Reactor T (°C) P (bar) Conv.a (%) Product selectivity Ref.
a Conversion.
1 Pt/Al2O3 Continuous 300 1.4 14 Phenol (65 mol%) 25
o-Cresol (15)
Benzene (4.4)
Cyclohexanone (3.3)
2,6-Dimethylphenol (2.4)
2-Methylanisole (0.77)
p-Cresol (0.22)
4-Methylanisole (0.13)
2 NiCu/Al2O3 Continuous 300 10 78.6 Cyclohexane (24.3 mol%) 79
Benzene (59.9)
Methylcyclohexane (8.8)
Toluene (2.9)
Cyclohexanol (2.3)
Cyclohexanone (1.8)
3 CoMoS/Al2O3 Batch 250 34.5 — Phenol, benzene and cyclohexane (∼82 wt%) 80
o-Cresol, toluene and methylcyclohexane (∼18)
4 CoMoS/Al2O3 Batch 300 50 58 Benzene (10 mol% of product) 50
Phenol (64)
Methylated products (10)
Cyclohexene, cyclohexanol, methylcyclohexane, and methylcyclohexene (3)
5 Rh/SiO2 Continuous 300 10 53.4 Aliphatic/aromatic products molar ratio (0.22) 28
6 Rh/CeO2 Continuous 300 10 100 Aliphatic/aromatic products molar ratio (0.30)  
7 Rh/ZrO2 Continuous 300 10 99.6 Aliphatic/aromatic products molar ratio (0.88)  
8 Rh/CoSiO3 Continuous 300 10 82 Aliphatic/aromatic products molar ratio (0.81)  
9 RhCo/Al2O3 Continuous 300 10 98 Aliphatic/aromatic products molar ratio (1.68)  
10 RhCo/SiO2 Continuous 300 10 99 Aliphatic/aromatic products molar ratio (0.30)  
11 Co/SiO2 Continuous 300 10 10.3 Aliphatic/aromatic products molar ratio (0.03)  
12 Ni/SiO2 Continuous 300 10 92.8 Aliphatic/aromatic products molar ratio (1.44)  
13 Ni/Cr2O3 Continuous 300 10 90.2 Aliphatic/aromatic products molar ratio (1.53)  
14 Ni/Al2O3 Continuous 300 10 80 Aliphatic/aromatic products molar ratio (3.02)  
15 Ni/ZrO2 Continuous 300 10 26 Aliphatic/aromatic products molar ratio (0.12)  
16 NiCu/Al2O3 Continuous 300 10 99.6 Aliphatic/aromatic products molar ratio (4.81)  
17 NiCu/ZrO2 Continuous 300 10 63.5 Aliphatic/aromatic products molar ratio (0.20)  
18 CoMoS/Al2O3 Continuous 300 10   Phenol (41 mol%)  
Benzene (23)
Cyclohexane (5)
Cresol isomers (30)
19 CoMoS/Al2O3 Continuous 250 15 88.2 Phenol (48.4%) 41
Benzene (7.5)
Toluene (4.0)
o-Cresol (24.7)
o-Methylanisole (1.4)
2,6-Xylenol (10.8)
20 CoMoS/Al2O3 Continuous 300 15 96.8 Phenol (40.8)  
Benzene (22.8)
Cyclohexane and cyclohexene (4.3)
Toluene (9.7)
o-Cresol (12.2)
2,6-Xylenol (3.7)
21 MoP/SiO2 Continuous 300 15 17.8 Phenol (∼75 mol%) 30
Cyclohexane (∼5)
Benzene (∼4.5)
22 Ni2P/SiO2 Continuous 300 15 76.8 Phenol (∼5)  
Cyclohexane (∼90)
Benzene (∼5)
23 NiMoP/SiO2 Continuous 300 15 28.6 Phenol (∼3)  
Cyclohexane (∼81)
Benzene (∼16%)
24 Pt/HBeta (1 wt% Pt) Continuous 400 1 ∼99 Yield of product (%): phenol (0.1), cresols (0.2), C1–2 (6), C3 (0.4), C4–9 (1.6), benzene (51.2), toluene (27.6), xylenes (10.6), C9+ aromatics (1.7) 81
25 CoMoWS/SBA-15 Continuous 310 30 38 Phenol (60.1%) 78
o-Cresol (25.9)
o-Xylenol (11.2)
o-Methylanisole (1.7)
Benzene (1.0)
26 CoMoWS/1 wt% P/SBA-15 Continuous 310 30 38 Phenol (62.7)  
o-Cresol (24.4)
o-Xylenol (10.7)
o-Methylanisole (1.7)
Benzene (0.5)
27 CoMoWS/SBA-16 Continuous 310 30 38 Phenol (60.8)  
o-Cresol (25.1)
o-Xylenol (10.8)
o-Methylanisole (1.8)
Benzene (1.5)
28 CoMoWS/1 wt% P/SBA-16 Continuous 310 30 38 Phenol (65.9)  
o-Cresol (22)
o-Xylenol (8.4)
o-Methylanisole (2.3)
Benzene (1.4)

3. Effective parameters on reaction selectivity of catalytic HDO of phenol, cresol, guaiacol and anisole

3.1. Catalyst promoters

Use of multi-component catalysts seems to be highly useful for selective hydrodeoxygenation of phenols to high-value hydrocarbons. Due to the high activity of transition metals in reactions including hydrogen, several elements of this metal group have been vastly used as catalyst for hydrodeoxygenation of phenolic compounds. Supported or unsupported transition metals have been used as both hydrogenation catalyst and promoter in HDO process. The addition of promoters to metal catalysts has been studied in order to determine their effects on catalytic activity and reaction selectivity. In HDO of anisole over bimetallic NiCu catalyst, the ratio of aliphatic to aromatic hydrocarbons in product stream was higher than that over mono-Ni catalyst demonstrating that the bimetallic active sites in NiCu/Al2O3 were more efficient for aromatic ring hydrogenation compared to single Ni active sites in Ni/Al2O3.79 Gutierrez et al.33 showed that combining Rh with Pt or Pd noble metals enhanced catalytic activity and guaiacol conversion due to the efficient interaction between these metals leading to higher hydrogen adsorption which facilitates hydrogenation process; conversion of about 10% on monometallic catalyst was enhanced to 32.7% on RhPd and 98.7% on RhPt (see Table 3, entries 12–17). However, the combination of Pd and Pt led to lower active surface area which caused lower guaiacol conversion and catalytic activity compared to individual Pd and Pt catalysts. Promoters affect reaction selectivity through the influence on physical (i.e. surface area, pore volume and particle size) and chemical (i.e. type of active sites) properties of catalyst. The influence of different promoters on the performance of HDO catalysts is illustrated in the following subsections.
3.1.1. Effect of promoters on catalyst structure. A major effect of promoters on catalyst is change of catalyst structure. For instance, introduction of promoter Co into Ni–Mo–B destroyed the shell structure of catalyst and caused more uniform and smaller particles.38 Study of the promoting effect of Mo in Co–Mo–B amorphous catalyst displayed that the increase of Mo content decreased catalyst surface area due to small BET surface area of Mo oxides.21 Meanwhile, it was shown that lanthanum (La) addition in an appropriate amount increased the amorphous degree of Ni–W–B catalyst due to enhanced interaction between hydrogenation active site of Ni0 and B0 which results in agglomeration of catalyst particles.37 However, excess content of La reduced the amorphous degree since large atom of La in catalyst acts as dispersing agent and could intercept particle agglomeration. In HDO of phenol over Ni–W–B, free hydrogen activated on Ni site was more involved in benzene ring hydrogenation of phenol molecule activated on WO3 Brønsted acid site than cleavage of Caromatic–OH bond of phenol. By addition of La into catalyst composition, this mechanism was not changed, phenol conversion decreased, deoxygenation rate increased and cyclohexanol, cyclohexene and cyclohexane were the main products.

Yoosuk et al. studied the effect of the addition of Ni35 and Co56 promoters on the activity of unsupported sulfided molybdenum amorphous catalyst in HDO of phenol. Nickel and cobalt addition leads to the formation of Ni–Mo–S and Co–Mo–S structures with weaker metal–sulfur bond strength in comparison with the Mo–S catalyst without promoter. Hence, the elimination of sulfur atom which leads to formation of sulfur vacancies (active sites) on the edge of MoS2 slabs can be easily occurred in amorphous NiMoS and CoMoS. Moreover, promoted catalysts had better hydrogen activation capability. However, Ni and Co differently alter catalyst properties and product selectivities. Ni addition into MoS2 catalyst led to an increase in phenol conversion.35 In the presence of Ni promoter, cyclohexane was the selective product and HDO process showed a strong favor toward HYD route. Ni addition can also form metallic-like brim sites (fully sulfided metallic edges and active sites for HYD reaction pathway) next to MoS2 slabs. Reduction of surface area, pore volume, number of large pores and number of layers in stacks are typical changes in amorphous MoS2 catalyst structure due to addition of Ni promoter. Nickel metal also changes catalyst morphology to crystalline structure. Long slabs of MoS2 became smaller and more curved by Ni addition due to the formation of smaller particles. Edge planes and basal planes of MoS2 layers are two models of active sites in Mo sulfide catalysts.56,82–85 Formation of large particles in catalyst prevents the prevalence of edge planes available for exposure to reactants. Also, Mo sulfide catalyst with bent basal planes have higher catalytic activity than that with flat basal planes. In fact, formation of small particles and predominance of bent basal planes in MoS2 layers are the positive effects of Ni promoter. However, cobalt metal in MoS2 catalyst acted as inhibitor of formation of crystalline MoS2.56 By adding cobalt as promoter, phenol conversion strongly increased and reaction selectivity shifted toward DDO route (see Table 1, entries 59 and 60). Addition of Co as promoter into amorphous MoS2 decreased catalyst surface area and pore volume and also led to shorter slabs in the stacks of catalyst. In catalytic HDO of guaiacol over molybdenum sulfide catalyst, addition of cobalt promoter led to remarkable increase in guaiacol conversion, catalytic activity and final oxygen-free products.23 HDO reaction was proceeded through DME and both DME and DMO routes over non-promoted and promoted molybdenum sulfide catalysts, respectively. On MoS2 catalyst, catechol was formed as the primary product followed by formation of phenol, methylcyclopentane, benzene and cyclohexane (as the major product). In the case of using CoMoS, phenol and catechol were the primary products and benzene (as the main product), cyclohexane, cyclohexene and toluene were produced as final products. Higher acidity of MoS2 compared to CoMoS resulted in greater amount of heavy methyl-substitution compounds production which harden the deoxygenation.

3.1.2. Effect of promoters on catalyst active site. Promoters are influential on catalytic activity and reaction selectivity through affecting type and dispersion of catalyst active sites. For instance, large radius atom of Mo acts as dispersing agent in molybdenum-promoted Co–Mo–B catalyst.21 Higher content of Mo in initial material decreases Co diffusion and inhibits agglomeration resulting in increase in the Co content of catalyst surface. High content of Co on catalyst surface is beneficial to adsorb enough hydrogen and supply sufficient free hydrogen while high content of MoO2 on catalyst surface is beneficial to provide enough Brønsted acid sites to activate oxy-groups.

In guaiacol HDO, incorporation of Co into unsupported molybdenum nitride catalyst (Mo2N) enhanced the selectivity of benzene, cyclohexane and cyclohexene as deoxygenated products.86 Cobalt addition into this catalyst resulted in formation of γ-Mo2N and Co3Mo3N active phases which lead to increase of deoxygenation activity. Higher production of oxygen-free compounds over cobalt modified catalyst is caused by Co3Mo3N active sites. Unlike unsupported Mo2N which was promoted by addition of cobalt, Co incorporation reduced catalytic activity of molybdenum nitride catalysts supported on alumina, mesoporous SBA-15 silica and activated carbon (AC) and did not cause any positive effect on HDO of guaiacol.20,87 The reason of this negative effect is that the Co3Mo3N active sites of deoxygenation were not completely generated on supported CoMoN catalysts and in the case of using CoMoN/AC, the concentration of cobalt in the internal part of catalyst was higher than that in the external part of support. High quantity of Co on surface of catalyst can enhance the number of active sites while high concentration of Co in the pores of catalyst reduces the accessibility of reactants to Co and also causes a restriction in diffusivity of reactants into the inner part of support. Incorporation of copper into nickel oxide catalysts supported on Al2O3, CeO2–ZrO2, SiO2–ZrO2–La2O3 and SiO2 facilitated the reduction of nickel oxide phases and increased catalytic activity due to decrease of coke formation rate.28,29,32 Copper promoter addition into Ni/Al2O3 catalyst reduced the formation of Ni spinels (NiAl2O4) which cause to decrease the rate of formation of active nickel metal (Ni0).79

Addition of promoter into catalyst composition in an optimum value have positive effect on HDO reaction rate while, an excess amount of promoter covers HDO active sites leading to decrease in catalytic activity. Ni addition into MoS2 catalyst led to an increase in phenol conversion and maximum conversion of 96.2% was achieved at Ni/(Mo + Ni) molar ratio of 0.3.35 At molar ratios above 0.3, a bulk phase of NiS was formed which covers the active part of promoted Mo sites and reduces phenol conversion. In HDO of phenol over Ni–Mo–B catalyst, it was revealed that the selectivity toward cyclohexane was increased by the addition of Co till an optimum level due to enhancement of hydrogenation–dehydration (see Table 1, entries 33–36).38 By further addition of Co, selectivity toward cyclohexane decreased since Ni active sites are covered by excess cobalt oxide leading to suppression of hydrogen adsorption on catalyst surface. In another work held by Song et al.88 for catalytic HDO of m-cresol over Pt/γ-Al2O3, it was shown that product distribution was strongly dependant on the amount of Pt loading (0.05–1.70 wt%). Direct deoxygenation of m-cresol to toluene was improved at high Pt loadings due to sufficient amount of metal active sites in catalyst structure. However, at low loadings of Pt and higher ratio of acid to metal sites, selectivity toward cracking reactions was favorably enhanced compared to hydrogenolysis mechanism.

In addition to metal promoters, non-metal elements could also be considered as promoters of HDO catalysts. Transition metal phosphide catalysts are another type of catalysts which are efficient in hydroprocessing.89,90 Nickel-, iron-, molybdenum-, cobalt- and tungsten-phosphide catalysts supported on neutral silica were applied for HDO of guaiacol in a packed bed reactor under 300 °C and atmospheric pressure.31 Metal sites dispersion and catalytic activity of these phosphide catalysts decreased in the order MoP/SiO2 > Ni2P/SiO2 > WP/SiO2 > Fe2P/SiO2 > Co2P/SiO2 and Ni2P/SiO2 > Co2P/SiO2 > Fe2P/SiO2 > WP/SiO2 > MoP/SiO2, respectively. The dominant pathway in guaiacol HDO over the transition metal phosphides was catechol formation via demethylation of guaiacol followed by phenol production via removing the –OH functional group of catechol and subsequently benzene formation by direct hydrogenolysis of phenol. Benzene and phenol were the main products of guaiacol HDO over phosphide catalysts. High capability of phosphide in production of phenol instead of catechol leads to low coke formation over phosphide-containing catalysts because the potential of catechol for coke generation is higher than phenol. Meanwhile, reaction selectivity of anisole HDO over Ni2P/SiO2, MoP/SiO2 and NiMoP/SiO2 catalysts at 300 °C and 15 bar hydrogen pressure was toward demethylation, hydrogenolysis and hydrogenation resulting in phenol, benzene and cyclohexane as the main products (see Table 4, entries 21–23).30 HDO activity of catalysts decreased in the order Ni2P/SiO2 > NiMoP/SiO2 > MoP/SiO2. High active Ni and Mo metal sites of these phosphide catalysts acted as Lewis acid sites. These bifunctional active sites participated in hydrogenolysis/hydrogenation and demethylation reactions through their roles as metal sites and acid sites, respectively. The high active Ni site was more effective than Mo site, and both these sites were more active than P–OH group as Brønsted acid site. Furthermore, Yang et al.16 studied the promoting effect of phosphorus in sulfided CoMo/MgO catalyst for HDO of phenol (see Table 1, entries 43 and 44). Phenol hydrotreatment over sulfided CoMo/MgO catalyst for 60 min led to a liquid product consisting 83.3% phenol, 3.2% cyclohexylbenzene, 7.5% cyclohexylphenol and a trace amount of benzene (0.01%). The mechanism suggested for this reaction is depicted in Fig. 8. To enhance the activity of CoMo/MgO, phosphorus was added into catalyst structure promoting the formation of MoS2 active site. 60 min treatment of phenol over sulfided CoMoP/MgO at 350 °C resulted in a liquid product containing 64.8% phenol, 6.4% cyclohexylbenzene, 13% cyclohexylphenol and 13.2% benzene. Decrease of coke formation, increase of Mo dispersion and also formation of new Brønsted and Lewis acid sites on surface of catalyst were the positive effects achieved by addition of phosphorus into CoMo/MgO catalyst.


image file: c5ra22137d-f8.tif
Fig. 8 Phenol hydrotreating mechanism for formation of cyclohexylbenzene, cyclohexylphenol and benzene over CoMo/MgO.16

3.2. Catalyst support

Selection of a suitable support is an important catalyst design parameter in HDO process since materials used as catalyst support play a key role in product selectivity of HDO reaction.66,91 A supported catalyst allows the catalytically active phase to be spread over the large internal surface of the support material. Heterogeneous catalysts are usually supported by reactive and neutral materials to have higher metal dispersion, catalytic activity and structural stability under harsh operating conditions of HDO process. In HDO of guaiacol, alumina- and SBA-15 silica-supported Mo nitride catalyst was shown to have higher activity than unsupported Mo nitride catalyst due to the improvement of particle dispersion and electronic interaction as well as higher accessibility of active sites to reactants caused by the porosity of supports.20 Supports with different properties such as acidity–basicity, pore structure and surface area may differently change the catalyst activity. Different performance of Pt supported on γ-Al2O3, SiO2 and HBeta in HDO of m-cresol revealed that pore structure and acid density of support are highly effective on catalytic activity; HBeta zeolite with higher density of acid sites and microporous structure led to lower activity.92 High density of acid sites favors cracking mechanisms and coke formation. Meanwhile, micropores are easily blocked by coke deposits leading to limited access of reactants to the active sites. Apart from catalyst activity, reaction pathway could also be affected by catalyst support. For instance, CoMoS catalyst supported on different materials of Al2O3, TiO2 and ZrO2 exhibited different reaction pathways in guaiacol HDO (Fig. 9).34 Demethylation and methylation, demethylation and hydrogenation as well as demethoxylation and direct deoxygenation were the dominant reaction pathways observed in guaiacol HDO over CoMoS/Al2O3, CoMoS/TiO2 and CoMoS/ZrO2, respectively. Alumina, activated carbon, silica, zirconia and zeolites are some examples of supports which have been used for the preparation of catalysts used in HDO of phenol, cresol, guaiacol and anisole. The influence of different catalyst supports on catalytic performance and reaction selectivity of HDO of phenols is reviewed in this section.
image file: c5ra22137d-f9.tif
Fig. 9 Principal reaction network for guaiacol HDO occurring on (a) CoMoS/Al2O3; (b) CoMoS/TiO2; (c) CoMoS/ZrO2.34

The effect of support nature on product selectivity was investigated using Pd catalyst (1–10 wt%) supported on mesoporous CeO2 and ZrO2 for HDO of phenol in a continuous fixed bed reactor at atmospheric pressure.57 At 180 °C and about 80% phenol conversion, the product obtained over Pd/CeO2 was a mixture of cyclohexanone (∼50%), cyclohexanol (∼35%) and cyclohexane (∼15%), while cyclohexanone (more than 90%) was the only remarkable compound produced by phenol HDO over Pd/ZrO2 (see Table 1, entries 61, 63, 67 and 68). This difference in product selectivity between Pd/ZrO2 and Pd/CeO2 is due to different adsorption positions of phenol on these catalysts caused by different acid–base properties of their supports. Over Pd/ZrO2 and Pd/CeO2, phenol is adsorbed on catalyst surface via non-planar and co-planar positions, respectively (Fig. 10). Phenol is adsorbed on basic sites via non-planar position resulting in high selectivity toward cyclohexanone, while acidic sites adsorb phenol molecule through co-planar mode producing more cyclohexanol and cyclohexane. In a study held by Ruiz et al.,74 the basic nature of zirconia was changed to acidic through sulfidation of ZrO2. It was observed that ReS2/ZrO2-sulfated catalyst had better dispersion of active sites compared to ReS2/ZrO2 whilst, guaiacol HDO reaction rate over the latter catalyst was higher than that on the former one. Higher acidic strength resulted from the presence of sulfate groups in sulfated support led to better dispersion of metal in ReS2/ZrO2-sulfated. The high activity of ReS2/ZrO2 is related to its structure; ReS2/ZrO2-sulfated had monolayer patches of ReS2 on support while, ReS2/ZrO2 had lamellar structure of sulfided Re with higher thickness. High thickness of ReS2 layers provides higher sulfur vacancies as catalytic active sites. In guaiacol HDO at 300 °C, 50 bar and over the two Re-based catalysts, phenol and catechol were the main products whereas, deoxygenated products such as cyclohexane, cyclohexene, hexane and benzene were formed in very low content.


image file: c5ra22137d-f10.tif
Fig. 10 Non-planar adsorption of phenol producing cyclohexanone (a); co-planar adsorption of phenol producing cyclohexanol (b).57

Alumina has been widely used as catalyst or support in industrial catalytic reactions.93 Intrinsic acidity of alumina support could lead to high catalytic activity through its role in enhancing the dehydration reaction in hydrodeoxygenation process. The overall HDO reaction of m-cresol over Pt/Al2O3 catalyst is proceeded through hydrogenation and dehydration routes catalyzed by metal and acid functions, respectively (Fig. 11).61 Toluene and methylcyclohexane were the main oxygen-free products formed over Pt catalyst supported on Al2O3. It was reported that alumina treatment with base K2CO3 and acid NH4F changes the reaction activity of Pt/Al2O3 catalyst (see Table 2, entries 27–35). Basic treatment of alumina leads to the deactivation of dehydration active sites due to the blockage of these phases caused by replacement of H+ of support Brønsted acid sites by K+. Increase of the acidity of support enhances the HDO rate due to the improvement of dehydration reaction. However, high acidity of alumina leads to high production of coke which deactivates catalyst rapidly. In HDO of guaiacol over Mo nitride catalysts supported on Al2O3, DME route was the major pathway due to high acidic properties of alumina and catechol resulted from DME route is a main intermediate compound which causes coke production.20 Meanwhile, it was reported by Bui et al.23,34 that in HDO of guaiacol over MoS2/Al2O3 and CoMoS/Al2O3, the acidic property of alumina results in production of methyl-substituted compounds which lead to difficult deoxygenation (see Table 3, entries 23–28).


image file: c5ra22137d-f11.tif
Fig. 11 Proposed reaction network for m-cresol conversion over Pt/Al2O3.61

Mesostructure silica is considered as a highly advantageous support for several types of functional groups.94 Uniform channels, large pore size, high surface area and high thermal stability are great characteristics of silica materials as support.95 Large pore size of mesoporous silica enhances mass transfer of reactant molecules. Generally, the interaction between active sites and silica support is effectively strong and can remarkably change catalytic property.96 In fact, weak metal-support interaction leads to sintering which results in catalyst deactivation. The mesoporous silica SBA-15 and SBA-16 are considered proper catalyst supports for HDO process because of their high ratio of surface/volume, changeable framework composition and high thermal stability.78 Hydrodeoxygenation of anisole on sulfided CoMoW catalyst and its modified form (by phosphate with concentrations of 0.5 or 1.0 wt%) supported on mesoporous silica SBA-15 and SBA-16 was carried out in a continuous packed bed reactor under 310 °C and 30 bar (see Table 4, entries 25–28).78 The capability of SBA-16 in dispersion of MoWS2 active sites was higher than SBA-15. Modifying the catalyst with phosphate can lead to the reduction of both total catalyst acidity (due to the high interaction between Mo(W) ions and Brønsted acid P–OH groups) and catalyst deactivation caused by coke and water.78,97 Catalytic activity of samples was decreased in the order CoMoW (0.5 wt% P)/SBA-16 > CoMoW/SBA-15 ≈ CoMoW/SBA-16 ≈ CoMoW (1.0 wt% P)/SBA-15 > CoMoW (0.5 wt% P)/SBA-15 > CoMoW (1.0 wt% P)/SBA-16. Except CoMoW (0.5 wt% P)/SBA-16 catalyst, phosphate incorporation led to the decrease of interaction between metal and support resulting in the formation of large metal sulfide species which can block support pores. The similar activity of non-modified catalysts supported on SBA-15 and SBA-16 illustrated that the morphology of these supports is not influential on catalytic activity.78 Splitting of Cmethyl–O bond occurs on both metal sulfides and acid sites of support. The acidic function of sulfided CoMoW/SBA-15(16) catalysts was more effective than their metallic function. Product selectivity in this process indicates two reaction routes: anisole demethylation which leads to phenol and methyl transfer to benzene ring which leads to o-cresol and o-xylenol formation. In a comparison between alumina and silica supported catalysts, it was reported by Foster et al.61 that activity of Pt/SiO2 in HDO of m-cresol is higher than that of Pt/Al2O3 with the same product selectivity. Also in HDO of guaiacol over MoN/SBA-15 and MoN/Al2O3 catalysts, DMO and DME were the major pathways on mesoporous silica- and alumina-supported catalysts, respectively.20 Density of acid sites in alumina supported catalyst was four times higher than that supported on SBA-15 silica leading to higher coke formation. Although the conversion of guaiacol on MoN/SBA-15 was lower than MoN/Al2O3, the lower amount of hydrogen utilization and coke formation were the two benefits of Mo nitride catalyst supported on mesoporous silica.

Solid acid zeolites are another type of support materials used for HDO catalysts. Indeed, zeolites could be used as both catalyst and support which participate in acid-catalyzed reactions and affect reaction selectivity. A mixture of Pd/C (hydrogenation function) and HZSM-5 as Brønsted solid acid (dehydration function) was shown to be a high potential catalyst for hydrodeoxygenation of phenol;55 at 200 °C, 50 bar and reaction time of 2 h, complete conversion of phenol led to 100 C% cycloalkanes selectivity. Meanwhile, catalytic transformation of phenol over Ni/HZSM-5 was carried out through hydrogenation and dehydration mechanisms proceeded over Ni metal and HZSM-5 support, respectively.98 In HDO of anisole in a packed bed reactor at 400 °C and atmospheric pressure using Pt noble metal catalyst supported on HBeta zeolite, Brønsted acid sites of HBeta catalyzed transalkylation reaction leading to phenol, cresols and xylenols as major products.81 Demethylation, hydrodeoxygenation and hydrogenation routes were sequentially catalyzed by Pt metal sites of catalyst to produce phenol, benzene and cyclohexane. A comparison between bifunctional Pt/HBeta catalyst and monofunctional HBeta and Pt/SiO2 catalysts exhibited that catalytic activity of Pt/HBeta catalyst in both transalkylation and HDO routes was higher and the amount of hydrogen used by this catalyst was lower (Fig. 12). Besides, Pt supported on acidic mesoporous ZSM-5 zeolite exhibited remarkable capability in hydrodeoxygenation of mixed isomers of cresol (above 93% deoxygenation degree).24 Gallium metal supported on HBeta and HZSM-5 zeolites was used for catalyzing HDO of m-cresol.62 At similar gallium loading, HDO activity of Ga/HBeta was higher than that of Ga/HZSM-5 (see Table 2, entries 36–38). Different activity of the two zeolites is related to the difference in their pore size. The pore size of beta zeolite is larger than that of ZSM-5 causing higher diffusion of large molecules of m-cresol into beta zeolite. In contrast to HBeta and HZSM-5, SiO2 is not a suitable support to stabilize gallium species and its low acidity leads to lower activity in HDO of m-cresol. Due to the easy adsorption of m-cresol on the surface of catalysts supported on zeolites, a surface pool of oxygen-containing compounds is probably formed in these catalysts.62,99 The surface pool of oxygen-containing compounds is an intermediate which can be converted to phenol, heavy oxygen-containing compounds, benzene, toluene, xylene and light C2–C6 hydrocarbons.


image file: c5ra22137d-f12.tif
Fig. 12 Proposed reaction pathways of anisole HDO over (a) HBeta; (b) 1% Pt/SiO2; (c) 1% Pt/HBeta.81

Activated carbon could be a suitable choice as catalyst support for hydrodeoxygenation of oxygen-containing compounds due to the low activity of activated carbon for coke generation. Centeno et al.66 indicated that using low acidic supports such as activated carbon instead of alumina could promote the activity of CoMoS catalyst. Large surface area of activated carbon in comparison with alumina, silica and amorphous-silica-alumina (ASA) is considered as a positive property for high dispersion of metal active sites.27 Besides, Nikulshin et al.91 showed that in guaiacol HDO, catalytic activity of CoMoS supported on carbon-coated alumina (CCA) was higher than that supported on alumina. Lower acidity of carbon-coated alumina, lower CoMoS phase size and higher CoMoS content caused enhanced activity of CoMoS/CCA compared to CoMoS/alumina. Selective hydrodeoxygenation of phenol to cyclohexane was reported to be similar for all Pd, Pt, Rh and Ru catalysts supported on carbon. Using Pd/C in neutral water at 200 °C for 0.5 h, phenol conversion and cyclohexanol selectivity were reported to be 100% and 98 C%, respectively. While in the same reaction conditions, Pd supported on Al2O3, SiO2 and ASA led to about 90% phenol conversion, and selectivities toward cyclohexanol and cyclohexanone were approximately 70 and 30 C%, respectively. Al2O3, SiO2 and ASA as Lewis acid supports have the potential of stabilizing cyclohexanone intermediate.27,100 Stabilization of cyclohexanone production suppresses further hydrogenation and cyclohexane production. High stability of carbon in non-neutral solutions compared to Al2O3, SiO2 and ASA leads to enhanced performance of carbon supported catalysts in basic and acidic solutions compared to those supported on Al2O3, SiO2 and ASA.

3.3. Solvent

Catalytic activity and reaction selectivity of HDO of phenols can be promoted by type of reaction solvent. The same reaction carried out in different solvents gives a completely different performance in terms of activity and selectivity. Catalytic aqueous-phase hydrodeoxygenation of phenol over palladium supported on C, Al2O3, SiO2 and ASA was studied using different solvents such as H2O, NaOH–H2O, CH3COOH–H2O and H3PO4–H2O at 50 bar hydrogen pressure (see Table 1, entries 4–7 and 11–19).27 In the case of using carbon as catalyst support, alkaline solution resulted in the same phenol conversion and product selectivity as neutral water (phenol conversion 100% and cyclohexanol selectivity 98 C%). Whilst, acidic solutions changed product selectivity of Pd/C in favor of cyclohexane (75 C% in acetic acid–water solution with pH = 2.6 and 85 C% in phosphoric acid–water solution with pH = 2.1). Higher acidity of solvent leads to higher cyclohexane production. In the case of using Pd/Al2O3, Pd/SiO2 and Pd/ASA, catalytic activity was reduced by replacing neutral water with basic solution. The weak decomposition of amorphous inorganic oxides in alkaline solution and also aggregation of Pd particles leads to low content of Pd on the surface of these catalysts reducing hydrogenation activity. The results obtained from all catalysts show that acidic solvent is more suitable than neutral water for cyclohexane selectivity. Yan et al.101 and Panneman and Beenackers102 reported that water as solvent has negative effect on dehydration of cyclohexanol to cyclohexene.

Activity of noble metal catalysts (Pd, Pt and Ru) supported on C and Al2O3 as well as conventional catalysts of CoMo/Al2O3, NiMo/Al2O3 and NiW/Al2O3 was investigated in HDO of p-cresol using water and supercritical n-heptane as solvent (see Table 2, entries 1–15).36 In water solvent, Pt catalysts exhibited highest HDO activity and selectivity toward toluene and methylcyclohexane. While due to high sensitivity of conventional catalysts to water which blocks active sites of these catalysts, HDO was almost not occurred in water using CoMo, NiMo and NiW catalysts supported on Al2O3. Conversion of p-cresol over Pt, Pd and Ru noble metal catalysts was approximately complete and toluene selectivity was low in supercritical n-heptane. Using n-heptane as solvent, CoMo/Al2O3 exhibited the greatest activity and toluene selectivity (97%) among the three conventional catalysts. At HDO reaction conditions of 300 °C, 83 bar pressure and batch reaction time of 1 h, mass transfer limitation was observed in water. However, in supercritical n-heptane, both reactant and hydrogen gas are totally miscible and pore diffusion is highly improved. Therefore, there was no any mass transfer limitation in n-heptane. In fact, high hydrogen accessibility in supercritical n-heptane led to hydrogenation of benzene ring to 4-methylcyclohexanol while in water, low accessibility of hydrogen resulted in hydrogenolysis pathway and toluene production. Apart from the effect of solvent on hydrogen accessibility, solvent polarity can also affect product selectivity of p-cresol HDO. In a polar solvent like water, p-cresol prefers to be vertically adsorbed on catalyst surface with its hydroxyl group (Fig. 13, I). Therefore, the reaction proceeds by hydrogenolysis pathway for toluene production. On the other hand, in non-polar solvent such as n-heptane, p-cresol prefers to be adsorbed on catalyst surface with its benzene ring resulting in ring hydrogenation and 4-methylcyclohexanol production (Fig. 13, II).


image file: c5ra22137d-f13.tif
Fig. 13 Proposed reaction mechanisms for hydroprocessing of p-cresol in polar solvent (I) and non-polar solvent (II).36

3.4. Catalyst preparation procedure

The performance of catalysts in hydrodeoxygenation of phenolic compounds of bio-oil to high-value hydrocarbons is strongly dependent on the method of catalyst preparation. Method of catalyst preparation determines the homogeneity, metal dispersion, resistance toward sintering and structural properties of catalysts which are highly effective on catalytic activity. There are several different ways to prepare HDO catalysts such as impregnation, precipitation, sol–gel, etc. Selection of a suitable method for catalyst synthesis is considered as an important step in selective hydrodeoxygenation of phenolic compounds. The effect of two different preparation methods of impregnation and sol–gel on the performance of CeO2–ZrO2, SiO2–ZrO2–La2O3 and SiO2 supported nickel oxide catalysts in guaiacol HDO was studied by Bykova et al.29,32 The overall catalytic reaction of guaiacol HDO over nickel oxide catalysts is based on three main routes. (i) Formation of 1-methylcyclohexane-1,2-diol due to migration of methyl group to the aromatic ring and hydrogenation of aromatic ring. (ii) Phenol formation due to both demethylation and partial deoxygenation of guaiacol followed by hydrogenation of aromatic ring of phenol to cyclohexanone and cyclohexanol which are reduced to produce cyclohexane. (iii) Benzene formation through elimination of both methoxy and phenolic functional groups and further hydrogenation of benzene to cyclohexane. The reaction selectivity over NiCu/CeO2–ZrO2 and Ni/SiO2 catalysts prepared by impregnation was toward alicyclic C5–C7 hydrocarbons, cyclohexanol, cyclohexanone and 1-methylcyclohexane-1,2-diol. However, on these catalysts, cyclohexane and benzene formation was low and cyclohexanone was the main product. On the other hand, NiCu/SiO2, NiCu/ZrO2–SiO2–La2O3 and Ni/SiO2 catalysts prepared by sol–gel method were very selective for alicyclic hydrocarbons (especially cyclohexane) formation. Although, the conversion of guaiacol on all catalysts was higher than 80%, those prepared by sol–gel method had the highest deoxygenation degree (92–97%). Catalyst analysis exhibited that the catalysts prepared by sol–gel method were highly active due to the great content of fine particles of reduced Ni films on the surface of silicate structure. These catalysts had lamellar structure leading to high specific surface area of nickel active sites and in turn high catalytic activity.

Five different structures of MoS2 catalyst obtained by different preparation procedures were examined for HDO of phenol in a batch reactor under 320–370 °C and 28 bar H2:103 MoS2 derived by in situ decomposition of molybdenum naphthenate (MoNaph), MoS2 derived by in situ decomposition of ammonium heptamolybdate tetrahydrate (AHM), commercial crystalline MoS2, exfoliated MoS2 dispersed in water (TDM-W) and exfoliated MoS2 dispersed in decalin (TDM-D). In the whole range of reaction temperature (320–370 °C), catalytic activity decreased in the order AHM > TDM-D > MoNaph > crystalline > TDM-W. AHM derived MoS2 catalyst had larger BET surface area and smaller particle size (in terms of average stack height, slab length and number of stack layers). This means that MoS2 catalyst derived from AHM has high dispersion in reaction solution, and much more active sites are involved in HDO reaction. The HDO activity of MoNaph derived MoS2 catalyst was lower than that of AHM derived one since MoNaph was not completely decomposed during catalyst preparation. The main products of phenol HDO over AHM derived MoS2 catalyst were benzene, cyclohexylbenzene and 4-cyclohexylphenol indicating that MoS2 catalyst derived from AHM favored DDO reaction route. Furthermore, the activity of TDM-W is lower than that of TDM-D since TDM-W catalyst particles are conglomerated after being transferred from the polar solvent of water into the non-polar solvent of n-hexadecane at the beginning of HDO reaction.

To enhance the production of high-value hydrocarbons in HDO of phenols, it is important to develop a catalyst possessing large surface area. Larger catalyst surface area results in higher dispersion of catalytic active sites and in turn enhanced conversion and reaction rate. Apart from the selection of a suitable catalyst support with high surface area, one way to increase catalyst surface area is the use of proper additive in catalyst preparation step. Additives can affect catalyst structure by formation of new compounds in the catalyst. Use of citric acid as an additive for preparation of MoP was revealed to be effective on catalyst morphology and p-cresol HDO.60 Catalyst surface which was coarse without using citric acid became flat by addition of citric acid into catalyst precursors.60,104 Also by the use of citric acid, catalyst surface area was increased. Elemental analysis of MoP catalysts prepared with or without citric acid exhibited that citric acid addition leads to the presence of C as citrate compound. Carbon can improve catalyst structure by restricting metal crystallites accumulation and enhancing metal dispersion in catalyst. Generally, on MoP catalysts prepared with or without citric acid, toluene, methylcyclohexane and trace quantity of 1,3-dimethylcyclopentane were produced from hydrogenolysis, hydrogenation and isomerization of p-cresol, respectively. In the case of using MoP prepared with citric acid, direct hydrogenolysis pathway was dominant over hydrogenation pathway while in the case of using MoP prepared without citric acid, the rate of both pathways were approximately the same.

Type of metal precursor used for catalyst preparation is another preparation factor which affects catalytic activity. Different chemical composition of metal precursors leads to the presence of different elements in catalyst structure which can change the quality of active metal dispersion. The performance of Ru/C catalyst prepared from ruthenium precursors of Ru(NO)(NO3)3, RuCl3 and Ru(acac)3 in HDO of phenol was compared with that of commercial Ru/C (see Table 1, entries 23–30).8 At 250 °C, 100 bar H2 pressure and 4.3 h operation time, hydrogenation–dehydration was the only reaction mechanism and no benzene was detected in product. It was observed that the highest yield of cyclohexane produced through hydrogenation was obtained by the commercial Ru/C. Ru/C prepared from RuCl3 produced more cyclohexane compared to the catalyst prepared from Ru(acac)3. Hydrogenation activity of the catalyst made of Ru(NO)(NO3)3 was remarkably low probably due to nitrogen remained on catalyst. This result is in accordance with that obtained in hydrogenation of phenol over Pd/C prepared from Pd(NO3)2 as a nitrate precursor of palladium.105 Nitrogen had negative effect on catalyst performance due to poor distribution of metal active sites in catalyst. Nitrogen presence in catalyst precursor led to higher deposition of metal on catalyst external surface as compared to bulk phase resulting in poor distribution of metal active sites and in turn lower hydrogenation activity. On the other hand, the catalyst prepared from RuCl3 as Ru precursor led to high hydrogenation activity due to the presence of Cl− which participates positively in active site preparation.8

In addition, tungstic acid (H2WO4) and two heteropolyacids of phosphotungstic acid (H3PW12O40) and silicotungstic acid (H4SiW12O40) were used as precursors of tungsten for the preparation of oxide catalysts of Ni–W, Ni–W(P) and Ni–W(Si), respectively.52 The acidity of Ni–W oxide catalysts decreased in the order Ni–W > Ni–W(P) > Ni–W(Si). Although the potential of all Ni–W catalysts in phenol HDO was suitable, the catalysts prepared from heteropolyacids exhibited higher activity in HDO process. The dispersion of metal oxides on the surface of Ni–W(P) and Ni–W(Si) was better than that on the surface of Ni–W. Also, surface exposure of Ni as the active sites for hydrogenation was higher in Ni–W(P) and Ni–W(Si) catalysts compared to Ni–W meaning that the use of heteropolyacids as precursors of tungsten species resulted in more favorable properties in catalyst structure than the use of tungstic acid.

As metal precursor is effective on performance of HDO catalysts, chemical source of non-metal components of catalyst is also influential on catalytic activity and product selectivity of reaction. Chemical source of catalyst components can control the reaction selectivity through affecting catalyst structure and catalytic active phases. For instance, it was shown that type of materials utilized for nitridation of nitride metal catalysts and sulfidation of sulfide metal catalysts is mainly effective on catalytic activity and reaction selectivity of these types of HDO catalysts.74,86 Unsupported molybdenum nitride catalysts (Mo2N) nitridated by thermal conversion in NH3 (ammonolysis) and/or thermal conversion in mixture of N2/H2 (reduction/nitridation) were used in guaiacol transformation.86 The catalyst prepared by ammonolysis had higher surface area and reaction rate compared to that prepared through reduction/nitridation. The source of nitrogen applied for nitridation was effective on type of active phases in catalyst. γ-Mo2N which is highly active phase for hydrodeoxygenation was the dominant nitride phase in the catalyst nitrided by NH3. High concentration of γ-Mo2N as well as high atomic ratio of N/Mo were the reasons for high activity of catalyst nitrided through ammonolysis. On the other hand, nitridation of catalyst by the mixture of N2/H2 reduced catalyst activity due to the presence of β-Mo2N0.78, MoO2 and Mo phases. β-Mo2N0.78 is resulted from the transformation of γ-Mo2N leading to the decrease of hydrodeoxygenation active phase of γ-Mo2N. MoO2 and Mo metal lead to formation of molybdenum nitrides with low surface area. It should be mentioned that interaction between chemical source and catalyst support is also effective on catalytic performance of catalyst. For instance, nitridation of Mo catalysts supported on Al2O3 and mesoporous SBA-15 silica through ammonolysis was more effective on catalytic activity than reduction/nitridation method (see Table 3, entries 46–51),20 while in comparison with ammonolysis method, nitridation of Mo supported on activated carbon by reduction/nitridation method resulted in higher catalytic activity due to the better scattering of Mo oxynitrides.87 Also, sulfiding Re/ZrO2 catalyst using gas flow of H2S mixed with either N2 or H2 indicated that type of sulfidation mixture (H2/H2S and N2/H2S) influenced catalytic activity via its effect on catalyst structure.74 N2/H2S sulfidation mixture led to higher reaction rate due to the decomposition of lamellar structure of ReS2 into spherical shape. Meanwhile in HDO of guaiacol, N2/H2S sulfidation mixture led to higher content of active sites on support surface followed by an increase in both DMO reaction rate and phenol/catechol ratio in product (see Table 3, entries 52–54).

Sometimes, catalysts are reduced by reducing agents to become activated before the use in HDO process. Reduction time of catalyst in preparation step is also influential on catalytic activity of HDO catalysts. Transformation of guaiacol on Mo2N/Norit catalyst reduced by hydrogen gas at 400 °C for 4, 6 and 8 h illustrated that the reduction time of 6 h resulted in the maximum HDO activity.72 Increase of the reduction time from 4 h to 6 h promoted the elimination of surface oxygen functional groups and led to the enhancement of reactants accessibility to Mo2N species. Whilst, further increase of reduction time led to decrease of catalytic activity due to the appearance of NHx surface species as well as H and Mo unsaturated species resulted from the cleavage of Mo–N bond by H2. Unsaturated species of Mo react with oxygen present on support surface or on subsurface of Mo2N leading to high formation of Mo oxynitrides and decrease of catalytic activity.

3.5. Operating conditions

Temperature is highly effective on reaction selectivity of exothermic HDO by affecting the stability of intermediates and their consequent conversion. The difference in hydrogen availability on catalyst surface is the reason for different product selectivities obtained at different temperatures. At temperatures higher than optimum value, hydrogenation activity is decreased due to the reduction of hydrogen adsorption which is an exothermic reaction. High availability of hydrogen is suitable for hydrogenation reaction due to higher hydrogen consumption in hydrogenation pathway compared to hydrogenolysis route. Therefore at high temperature, more hydrogenolysis is taken place in comparison with hydrogenation. It was revealed that in HDO of phenol over sulfided catalysts, DDO and HYD routes are dominant reaction pathways occurred at high and low temperatures, respectively.56,103 Yang et al.16 showed that increase of reaction temperature up to 450 °C enhanced catalyst activity of CoMoP/MgO and benzene selectivity; phenol conversion over CoMoP/MgO and selectivity of benzene were increased from 2 and 1% at 300 °C to 89.8 and 64% at 450 °C, respectively. Ni–W oxide catalyst supported on activated carbon was used for HDO of phenol at four different temperatures (150, 200, 250 and 300 °C) and operating pressure of 15 bar (see Table 1, entries 49–52).52 At 150 °C, the activation of tungsten sites and adsorption of phenol was low and high yield of cyclohexanol was obtained. Cyclohexanol was also shown to be highly stable at low temperature of 150 °C in HDO of phenol over Pd/C combined with H3PO4 (see Table 1, entries 9 and 10).27 Increase of temperature in the range of 150–300 °C was shown to enhance the activity of Ni–W oxide catalyst.52 Besides, by increase of reaction temperature up to 250 °C, phenol conversion was increased and product selectivity was shifted toward cyclohexane. By increase of temperature from 250 to 300 °C, production of cyclohexane was decreased while production of benzene, methylcyclopentane and cyclohexene was increased indicating that at temperatures from 250 to 300 °C, hydrogenation activity was decreased over Ni–W catalyst.

Temperature dependence of catalytic activity and product selectivity of guaiacol HDO showed that increase of operating temperature (350–450 °C) promotes catalytic activity of Fe/SiO2 catalyst but has no effect on product selectivity.42 However, in HDO of guaiacol over NiCu/ZrO2–SiO2–La2O3 at 280, 320 and 360 °C, increase of reaction temperature led to the reduction of guaiacol conversion.32 High temperature provided high potential for polymerization of oxygen-containing organics. Therefore, coke formation was increased due to the coverage of catalyst active sites by large molecular compounds resulting in less conversion of guaiacol. It was reported that low and high temperatures were more selective toward oxygen-containing aliphatics and cyclohexane/benzene, respectively. In addition, dependency of product selectivity of guaiacol HDO over Rh/SiO2–Al2O3 catalyst to reaction temperature exhibited that by increasing temperature from room temperature to 250 °C, at 58–108 °C, guaiacol was completely hydrogenated and was mostly converted to 2-methoxycyclohexanol.70 Further heating the reaction system (to 250 °C) led to higher production of partially deoxygenated compounds such as cyclohexanol, cyclohexanone and fully deoxygenated compound of cyclohexane. Temperatures higher than 250 °C were suitable for production of fully deoxygenated compound of cyclohexane. HDO of guaiacol over Rh/ZrO2, PtRh/ZrO2 and PdRh/ZrO2 catalysts in a batch reactor at different temperatures (300, 350 and 400 °C) and 50 bar H2 pressure included hydrogenation of aromatic ring (yielding 2-methoxycyclohexanol and 2-methoxycyclohexanone) followed by demethoxylation and dehydroxylation producing cyclohexane.106 Increase of reaction temperature from 300 to 400 °C enhanced yield of cyclohexane. Besides, in HDO of m-cresol using MoO3 in a packed-bed reactor, it was shown that increase of reaction temperature form 300 to 400 °C reduced activity of MoO3 catalyst due to the change of active phase of MoO3 to inactive phase of MoO2.107 At 400 °C, complete conversion of MoO3 to MoO2 was occurred after 0.5 h reaction time and up to 80% of catalyst activity was reduced after 3 h. However, different behaviour was observed at lower temperature of 320 °C; MoO3 was changed to an oxycarbohydride phase (MoOxCyHz) with low amount of MoO2 after 1.5 h. It was mentioned that carburisation of MoO3 to MoOxCyHz stabilizes active phase of MoO3 and prevents from formation of inactive phase of MoO2.

Product distribution in HDO of phenolic compounds is a function of reaction temperature due to the temperature dependence of activation energies for different reaction routes taken place in HDO process. High operating temperature is needed for reaction mechanisms with high activation energy. In catalytic HDO of anisole over Pt/Al2O3, demethylation to phenol, demethoxylation to benzene and transalkylation to methylphenols were reported to be occurred at reaction temperatures in the range of 300–400 °C.108 It was shown that the activation energy for oxygen removing mechanism was lower than that for transalkylation mechanism. Therefore, selectivity toward deoxygenation was favored at lower operating temperatures.

In addition to temperature, pressure is another operating factor influential on selective HDO of phenols. High pressure of hydrogen is usually required to accelerate catalyst reduction and to enhance the efficiency of HDO process. In HDO of m-cresol over Pt/Al2O3 in a continuous fixed bed reactor at 260 °C, the increase of hydrogen pressure (0.5–1 bar) enhanced both m-cresol conversion (from 38 to 74%) and methylcyclohexane selectivity.61 The effect of hydrogen partial pressure (0–0.9 bar) on catalytic activity of Fe/SiO2 was investigated in atmospheric hydrodeoxygenation of guaiacol for selective production of benzene and toluene (products of complete HDO) instead of phenol (product of partial HDO).42 It was revealed that conversion of guaiacol in the atmosphere with no hydrogen and 0.2 bar H2 partial pressure was 30 and 70%, respectively. Guaiacol conversion was approximately constant in hydrogen partial pressure range of 0.2–0.9 bar. Increase of H2 partial pressure from 0.2 to 0.9 bar slightly increased the production of benzene and toluene. The maximum conversion of 74% and benzene and toluene production of 38 C% were achieved at H2 partial pressure 0.9 bar. In another work held by Saidi et al.,108 increase of reaction pressure in a fixed-bed tubular microflow reactor from 8 to 14 and 20 bar reduced conversion of anisole from 48 to 30 and 29%, respectively. However, pressure increase improved selectivity of HDO products. Selectivity toward benzene and phenol was increased, while selectivity toward methylphenols and hexamethylbenzene was decreased. Investigation of pressure influence on product distribution of HDO of phenols using mixed catalysts of Pd/C and ZSM-5 zeolite in a batch reactor showed that for hydrogen pressure in the range of 30–50 bar, HDO activity was almost remained unchanged, and selectivity of cycloalkane was 90% at reaction time of 0.5 h.55 For the pressures below 20 bar, yields of cycloalkane and cycloketone were 76 and 10%, respectively. By decrease of hydrogen pressure to 10 bar, reactant conversion was decreased to 80% and product selectivity toward cycloalkane and cycloketone was 12 and 75%, respectively. It was deduced that low hydrogen pressure stabilizes formation of cycloketone and hinders production of cycloalkane.

3.6. Co-feeding

Conventional hydrotreating catalysts are based on sulfided molybdenum in combination with a promoting metal such as cobalt or nickel. Typically, the catalysts are made in an inactive oxide form of active metals which have to be sulfided in order to reach their active forms. The catalysts are sulfided by the reaction of their active metals with hydrogen sulfide which can be produced from a reaction between the sulfur available in sulfiding agent and hydrogen (eqn (1)–(3)). A defined quantity of sulfur is required to maintain the catalysts in their active form. Since during the HDO reaction over sulfided catalysts, sulfur leaching from catalyst surface is happened, the addition of sulfiding agent during the reaction seems to be essential for maintaining the sulfided state of catalyst. H2S/H2 mixture gas or CS2 as a precursor for H2S are common sulfiding agents which have been utilized in HDO process. However, addition of sulfiding agent reduces catalytic activity and remarkably changes product selectivity. In fact, the adsorption of thiol group on catalyst surface leads to obstruction of the sites of hydrogen chemisorption decreasing the activity of catalyst for HDO reactions.109
 
MoO3 + 2H2S + H2 → MoS2 + 3H2O (1)
 
CoO + H2S → CoS + H2O (2)
 
3NiO + 2H2S + H2 → Ni3S2 + 3H2O (3)

The addition of sulfiding agent into HDO feed remarkably decreases initial activity of sulfided CoMo/Al2O3 catalyst but does not have clear positive or negative effect on catalyst stability.41,47 In the absence of sulfiding agent, sulfided CoMo/Al2O3 is much more selective toward benzene than cyclohexane and cyclohexene. However by addition of sulfiding agent, direct hydrogenolysis of phenol to benzene is strongly suppressed even in a very low amount of H2S. While, hydrogenation–hydrogenolysis conversion of phenol to alicyclic products is the same in the absence and presence of H2S (up to an optimum value) in feed stream. Therefore, selective inhibition of DDO reaction route by H2S addition can be considered as an effective way for enhanced production of alicyclic products. The selective inhibition of hydrogenolysis route by H2S illustrates that sulfided CoMo/Al2O3 contains at least two types of active sites which are different in sensitivity to sulfur.41 Direct hydrogenolysis reaction is activated by sulfur anion vacancies present in MoS2 which is the active site for HDO reactions in molybdenum sulfide catalysts; phenol is adsorbed on sulfur anion vacancies through its oxygen atom resulting in production of benzene and water. By addition of H2S, this route is inhibited as the result of competitive adsorption of phenol and H2S on sulfur anion vacancies. Higher electron affinity of hydrogenolysis active sites compared to hydrogenation sites leads to adsorption of H2S on hydrogenolysis sites and selective inhibition of hydrogenolysis route. However, addition of H2S inhibits both reaction routes of HYD and DDO in phenol hydrodeoxygenation over sulfided NiMo/Al2O3 catalyst.47 Meanwhile in HDO of guaiacol over sulfided Re/ZrO2, inhibitory effect of high concentration of CS2 sulfiding agent was the same as H2S leading to decrease of reaction rate and increase of catechol selectivity.74

Addition of sulfiding agent can have negative effect by being a cause for formation of sulfur containing byproducts. By addition of H2S, cyclohexanethiol is produced as a byproduct during phenol HDO over CoMo/Al2O3 and NiMo/Al2O3.47 Cyclohexanethiol can be produced through nucleophilic substitution reactions catalyzed by acid (Fig. 14). The existing mobile SH− groups on catalyst surface are able to act as nucleophiles and dispose catalysts for substitution reactions.47,110 The concentration of nucleophiles is a principal parameter in acid-catalyzed nucleophilic substitution reactions. Therefore, the production of thiol in the presence of sulfiding agent can be attributed to the increase in the number of SH− groups of catalyst surface caused by H2S addition.


image file: c5ra22137d-f14.tif
Fig. 14 Hydrogenation of cyclohexanol to form cyclohexene and cyclohexanethiol by elimination (E1 and E2) and nucleophilic substitution (SN1 and SN2) reactions, respectively.47

A reductant agent is usually mixed with oxygenates in order to facilitate the removal of oxygen. The reductant agent is decomposed into hydrogen molecule or prepares active hydrogen atoms to participate in hydrogenation/hydrogenolysis and hydrogen transfer reactions. The effect of co-feeding reductant agent of α-terpinene on catalytic activity and product selectivity was studied in HDO of guaiacol over vanadium oxide on alumina.71 Co-feeding α-terpinene with guaiacol led to easier oxygen elimination and higher yield of phenol as well as lower amount of coke generation. High tendency of V2O5/Al2O3 to phenol production is due to the oxophilic behavior of vanadium as well as its high potential to undergo changes of oxidation state. Vanadium is considered as a suitable catalyst for deoxygenation of the aromatic compounds containing oxygen atoms placed in 1,2-positions. Alumina support of vanadium oxide catalyst had impressive role in catechol production from guaiacol. In HDO of guaiacol co-fed with α-terpinene over V2O5/Al2O3, vanadium is reduced by α-terpinene and the organic compound of p-cymene is formed as the product of oxidation of α-terpinene (Fig. 15a). The reduced vanadium is bound to catechol via one of its oxygen atoms (Fig. 15b). The total scheme of guaiacol HDO (combination of Fig. 15a and b) is depicted in Fig. 15c.


image file: c5ra22137d-f15.tif
Fig. 15 p-Cymene formation due to reduction of vanadium by α-terpinene (a); the adsorption of oxygen atom of catechol by V(III) (b); the total scheme of guaiacol HDO on V2O5/Al2O3 catalyst (c).71

Co-feeding a hydrodesulfurization (HDS) feed with bio-based HDO feed leads to direct production of a partially bio-sourced fuel. In such systems, both HDS and HDO reactions are competitively proceeded. Straight run gas oil (SRGO, a residue in petroleum refineries) was co-fed with guaiacol in HDO process over CoMo catalyst supported on alumina.111,112 In a trickle bed reactor under 40 bar, temperature of below 320 °C and LHSV (liquid hourly space velocity) of 2 h−1, guaiacol was preferentially transformed into phenolic compounds and conversion of sulfur-containing molecules was inhibited.111 Intermediate phenolic products of guaiacol HDO was more adsorbed on hydrogenation and/or hydrogenolysis sites of catalyst compared to sulfur-containing compounds. In fact, at temperatures lower than 320 °C, the HDO feed acted as inhibitor for HDS reaction while HDS could be occurred after complete conversion of phenolic intermediates into deoxygenated compounds. Similarly, by co-feeding 4,6-dimethyldibenzothiophene (4,6-DMDBT) with guaiacol over ReS2/SiO2 and ReS2/Al2O3 catalysts, the adsorption of guaiacol molecules on catalytic active sites was more than 4,6-DMDBT molecules.113 Without co-feeding 4,6-DMDBT, ReS2/Al2O3 had higher total rate of guaiacol conversion due to stronger acidity of alumina and ReS2/SiO2 had higher HDO rate (larger amount of oxygen-free products) due to lower acidity of silica. By co-feeding 4,6-DMDBT with guaiacol, total rate of guaiacol conversion over ReS2/SiO2 was increased while HDO rate on this catalyst was decreased. On the other hand, both rates of guaiacol conversion and HDO were increased over ReS2/Al2O3 by 4,6-DMDBT addition. Adsorption of 4,6-DMDBT on ReS2/Al2O3 blocks Lewis acid sites of support which are active phases for demethylation route. Thus, less accessibility of support acid sites leads to lower selectivity toward DME pathway and catechol production as well as enhanced selective conversion of guaiacol to phenol via DMO followed by more hydrogenation of phenol to oxygen-free compounds.

4. Deactivation of HDO catalysts

In heterogeneous catalysis, loss of catalyst activity with time-on-stream is called deactivation. The knowledge about deactivation mechanisms and effective parameters on catalyst deactivation could help in designing more effective HDO process with high yield of desired products. Lower catalyst deactivation leads to longer catalyst lifetime and higher selectivity toward deoxygenated products.1,2,4,7–9 Catalyst stability is a key catalyst characteristic which determines its applicability in commercial scale. Generally, deactivation of HDO catalysts is caused by coking, poisoning, thermal degradation and sintering.

4.1. Deactivation caused by coke

The main challenge of HDO process is the rapid loss of catalyst activity due to formation of coke which is a competing reaction with production of aromatics and alicyclics. Polymerization and polycondensation are two main reactions resulting in coke generation which blocks catalytic active sites and reduces their accessibility for reactants.114 Coke deposition on catalyst surface is strongly dependant on type of reactants and intermediates. Unsaturated oxygenates such as aromatic fragments have high tendency to produce coke through polymerization on catalyst surface.115 Phenols are considered as a potential group of bio-oil compounds for coke production in HDO reaction.9,116,117 In comparison to phenol, cresol and anisole, guaiacol with two oxygen atoms bonded to benzene ring causes highest coke formation.1,2,4,7–9,19,61,80,117–119 Oxygenated compounds with more than one oxygen atom have high affinity for coke formation through polymerization mechanism.120 Catechol which is produced as an intermediate in HDO of guaiacol is also a strong agent for coke generation.19,20,31,42,61,71–73,119 In catalytic HDO of guaiacol over Ru/C and Pt/C, naphthalene which was produced from ring condensation mechanism and its larger condensed ring derivatives were reported as the main cause for rapid catalyst deactivation.121

In addition, deactivation of catalyst by coke is dependent on catalyst properties such as acidity. Generally, higher acidity of catalyst (including both Lewis and Brønsted acid sites) leads to higher coke formation.114 Lewis acid sites bind reactant species to catalyst surface, and Brønsted acid sites give proton to reactant molecules to form carbocations which are known to be coke precursor.115,122 In a study held by Zanuttini et al.92 for HDO of m-cresol over Pt/SiO2, Pt/Al2O3 and Pt/HBeta catalysts, Brønsted acid sites were reported to have the highest impact on catalyst deactivation by coking, and coke yield was shown to be reduced by decrease in density of Brønsted acid sites. Pt/SiO2 had lower total density of acid sites compared to Pt/Al2O3, but coke formation on Pt/SiO2 was higher than that over Pt/Al2O3. This was due to the presence of Brønsted acid sites in Pt/SiO2, while Pt/Al2O3 had only Lewis acid sites. In comparison to Pt/SiO2 and Pt/Al2O3, faster deactivation of Pt/HBeta was observed which was due to the high density of Brønsted acids of zeolite support. Meanwhile, based on literature reports, catalysts supported on alumina suffer from deactivation through coking owing to the presence of strong acid sites on this support. Enhanced HDO activity was obtained in guaiacol transformation over CoMoS supported on carbon-coated alumina compared to alumina.91 Coverage of alumina by carbon reduced total acidity of catalyst and resulted in lower coke formation. Eqn (4) was given to estimate degrees of catalyst deactivation in guaiacol conversion and HDO using a fixed bed microreactor:

 
image file: c5ra22137d-t1.tif(4)
where DdGua, DdHDO, xinitGua, xinitHDO, xsGua and xsHDO are catalyst deactivation degree in guaiacol conversion, catalyst deactivation degree in guaiacol HDO, initial guaiacol conversion, initial guaiacol HDO, steady-state guaiacol conversion and steady-state guaiacol HDO, respectively. Experimental data showed that degree of catalyst deactivation was a function of carbon amount used for coating the alumina support. DdGua and DdHDO for CoMo/Al2O3, CoMo/C/Al2O3 (2 wt% C) and CoMo/C/Al2O3 (5.6 wt% C) were 46, 33, 31% and 69, 19, 4%, respectively. These data indicate that increase of carbon content in catalyst support (or decrease of catalyst acidity) resulted in reduction of catalyst deactivation degree. This result is in agreement with that reported by Sun et al.123 who showed that substitution of alumina support with carbon enhanced catalyst stability in HDO of phenolic compounds since strong reaction of phenolic compounds and acidic sites of alumina leads to rapid deactivation of catalyst by formation of coke precursors.

Apart from catalyst properties, operating parameters such as temperature and pressure are influential on coke yield of HDO. Low hydrogen pressure and low reaction temperature favor the condensation/polymerization reactions rather than hydrogenation reaction. Indeed, hydrogen regenerates active oxygen vacancy sites and prevents from site blocking by carbonaceous species.124 Reaction temperature could be optimized for reduced coke yield. Increase of reaction temperature could result in lower coke yield due to hydrogenation of coke deposited on catalyst to stable products.125 However, temperatures higher than optimum value lead to high coke formation due to generation of hard coke with low volatility and solubility which is difficult to be desorbed from catalyst surface.28,122,126 In phenol HDO over Pd/HY–Al2O3, high coke formation and rapid catalyst deactivation was observed at temperatures higher than 250 °C.127 In catalytic HDO of phenol over Ni/Na–Y, effect of reaction temperature on degree of catalyst deactivation and initial conversion of phenol was quantitatively studied using an empirical model as follow:128

 
x = xinitial[thin space (1/6-em)]exp(−αΔt) (5)
where α (h−1), xinitial and t (h) refer to catalyst deactivation degree, initial conversion of phenol and reaction time, respectively. Increase of reaction temperature from 200 to 300 °C increased deactivation degree of Ni/Na–Y catalyst. Increase of reaction temperature led to higher pore blockage of catalyst due to higher amount of coke deposited on catalyst.

In addition to acid sites, metal active sites could also cause coke formation in HDO of phenolic compounds. However, deactivation rate of metal and acid sites of catalyst by coke deposition is different; in catalytic HDO of m-cresol over bifunctional catalysts of Pt/Al2O3 and Pt/HBeta, deactivation of acid sites was faster than that of metal sites.92 Deactivation degree of Fe/SiO2 catalyst in guaiacol HDO carried out in a fixed bed reactor was determined by fitting experimental data with an empirical model (eqn (6)).129

 
k = k0[thin space (1/6-em)]exp(αt) (6)
here, k (kmol s−1 kg−1), k0 (kmol s−1 kg−1), α (min−1) and t (min) are reaction rate, initial reaction rate, deactivation coefficient and time on stream, respectively. Deactivation coefficient values for guaiacol conversion, phenol transformation to benzene and cresol conversion to toluene (HDO mechanisms) were in the range of 4.6–5.03 × 10−3 min−1, while deactivation coefficient value for guaiacol transformation to cresol (transalkylation mechanism) was 7.29 × 10−3 min−1. Modification of structural properties of catalyst could enhance catalyst durability by decrease of coke formation over metal active sites. For instance, coke formation on metal Ni catalyst was remarkably suppressed by addition of Cu into catalyst structure leading to reformation of metal active phase by formation of NiCu alloy.130 Meanwhile, in a study held by Bykova et al.,114 catalyst deactivation through coke formation was decreased by addition of P and Mo to nickel based catalyst of NiCu/SiO2–ZrO2. The P and Mo additives changed the distance between Ni sites and thus hindered polymerization mechanism and coke formation.

4.2. Deactivation caused by poisoning

Catalyst poisoning caused by strong chemisorption of reactants, products or impurities on catalytic sites is considered as another main cause for catalyst deactivation during HDO reactions. In catalytic HDO of guaiacol over Ni/ZrO2, rapid deactivation of catalyst was occurred by co-feeding 0.05 wt% sulfur.131 Complete loss of activity was observed due to formation of inactive phase of nickel sulfide. Catalysts based on nickel have high potential to be deactivated by sulfur. Conventional hydrotreating catalysts such as NiMoS and CoMoS are rapidly deactivated by oxygenated compounds and water.132 Water which is used as solvent or produced through HDO reactions can lead to catalyst deactivation by being adsorbed on catalytic active sites.2,5,10,13–15,33,78,133 Water causes oxidation of active metals22,27,30,35–37 and agglomeration of particles and in turn loss of catalytic activity.134 It was reported by Mortensen et al.135 that catalyst deactivation rate by the water used as solvent is higher than that by the water produced from HDO reaction. In a study held by Olcese et al.136 for guaiacol HDO, oxidation of iron by phenols, water and carbon oxides was the main cause for deactivation of Fe supported on silica and activated carbon. In HDO of bio-oil carried out by Laurent and Delmon,137 it was shown that high content of water leads to deactivation of sulfided NiMo/Al2O3 catalyst through oxidizing nickel species. Meanwhile, Li et al.30 reported that Ni2P catalyst was deactivated by water through oxidization of active phase of phosphide to less active phases of phosphate or oxy-phosphide. It was also reported that water results in change of active phase of Mo2C to inactive phase of MoO2 in HDO of phenol135 and anisole.138 However, Prasomsri et al.124 pointed out that water addition did not increase deactivation rate of MoO3 catalyst used in HDO of phenolic compounds. MoO3 showed high tolerance to poisoning with water. Badawi et al.133 also reported that water was not the main cause for deactivation of Mo/Al2O3 catalyst in HDO of phenol, and addition of Co into catalyst structure decreased deactivating effect of water. Furthermore, in catalytic HDO of phenolic compounds over bimetallic catalyst of Pd/Fe2O3, Pd addition to Fe catalyst decreased Fe oxidation by water and catalyst deactivation rate due to the electronic interactions of Pd–Fe which facilitate the desorption of intermediates/products form catalyst active sites.139,140 Traditional hydrotreating catalysts supported on alumina are rapidly deactivated at high concentration of water due to the instability of alumina in contact with water.2,5,10,13–15,18,33 Water converts alumina to boehmite (AlO(OH)) leading to oxidation of metal catalysts and catalyst deactivation. Therefore, selection of a suitable support with high tolerance to water can increase the lifetime of HDO catalyst. For instance, use of hydrophobic activated carbon as support hinders catalyst deactivation caused by water.52,87 Moreover, the increase of water concentration can change product selectivity; according to chemical equilibrium of the reaction of cyclohexanol dehydration, the increase of water content decreases cyclohexanol dehydration and as a consequence leads to decrease of oxygen-free products.27–32 Loricera et al.78 pointed out that modifying the surface chemistry of CoMoW/SBA-16 with 0.5 wt% phosphate prevents catalyst deactivation by water and coke in HDO of anisole.

4.3. Deactivation caused by thermal degradation

Thermal degradation causes catalyst deactivation through sintering or chemical transformation of active phases. Morphological changes of catalyst due to sintering and agglomeration result in particle growth and reduction of catalytic activity.134 Loss of catalytic surface area caused by both crystallite growth of catalytic sites and pore breakdown on crystallites of active sites as well as loss of support area due to support breakdown generally occur by sintering phenomenon. In catalytic HDO of guaiacol over Pt, Pd, Rh and Ru supported on activated carbon, increase of reaction temperature from 275 to 325 °C resulted in decrease of catalyst surface area due to metal crystallite growth and collapse of support structure.121 Reduction of catalyst surface area results in lower catalytic activity since the amount of active sites available for reactant molecules is reduced. Average sizes of metal particles was increased due to sintering of catalysts; average sizes of Pt/C and Ru/C were increased from 2.40 ± 0.54 and 2.56 ± 0.47 nm to 2.67 ± 0.62 and 2.87 ± 0.63 nm, respectively. In addition to sintering, chemical transformation of active phases to inactive phases could be occurred as a result of thermal degradation of catalyst during HDO process. For instance, in HDO of lignin-derived oxygenates (phenol, m-cresol, anisole and guaiacol) over MoO3, increase of reaction temperature from 300 to 400 °C caused strong deactivation of catalyst due to the change of active MoO3 to inactive MoO2; at 400 °C and after 0.5 h reaction time, MoO3 was completely converted to MoO2.107 However, at temperatures below 350 °C, MoO3 was changed to oxycarbohydride-containing phases (MoOxCyHz) with minor impurity of MoO2. In fact, conversion of MoO3 to MoOxCyHz which is occurred via surface carburization in the presence of a carbon source like phenolic compounds leads to stabilization of MoO3 and reduction of MoO2 formation. First order deactivation model (eqn (7)) was used to determine catalyst deactivation degree by experimental data obtained from m-cresol HDO over MoO3 at different reaction temperatures.107
 
image file: c5ra22137d-t2.tif(7)
where x, kDeact (h−1) and t (h) refer to fractional conversion, first order deactivation rate constant and time, respectively. At 400 °C, rapid deactivation of catalyst was occurred and deactivation constants of 0.898 and 0.156 h−1 were achieved for reaction time regimes of 0–4 and 4–7 h, respectively. At reaction temperatures below 350 °C, catalyst deactivation was slow; deactivation constants at temperatures of 300, 320 and 350 °C were 0.048, 0.058 and 0.149 h−1, respectively.

4.4. Deactivation caused by desulfurization

In the case of using sulfided HDO catalysts, desulfurization also results in catalyst deactivation in addition to coke formation and water poisoning.2,5,10,13–15,114,141–143 Considering the low content of sulfur in bio-oil and in order to compensate sulfur leaching from catalyst surface and maintain sulfided state of catalyst, co-feeding a sulfiding agent seems to be necessary. Sulfiding agent in an optimum value can maintain catalytic activity of sulfided catalysts. However, at values higher than the optimum value, sulfur of sulfiding agent acts as an inhibitor of HDO reaction. Sulfur as competitor of oxygen-containing compounds is adsorbed on catalyst and deactivates the catalyst by decreasing the concentration of coordinatively unsaturated active sites (CUSs) of catalyst.41,144 By adsorption of sulfur-containing compounds on the CUSs, sulfur is present in final product composition. For sulfided HDO catalysts, very high temperature also leads to rapid catalyst deactivation caused by high concentration of water generation, large amount of sulfur loss and high yield of coke formation.43

5. Kinetic studies of HDO of phenols

Reaction kinetics is the study on quantification of rates of chemical processes. Rate of chemical reaction is the speed of change in concentrations or amounts of either reactants or products. Nature of reactants, catalyst type and operating conditions are the influential factors on rate of chemical reaction. Different studies have been conducted in order to explain HDO kinetics of bio-oil model compounds such as phenol, cresols, anisole and guaiacol. Kinetic study of HDO of bio-oil model compounds could be useful for finding catalysts effective for deoxygenation of bio-oil.45 Kinetically analysis of HDO reactions quantitatively illustrates the influence of experimental conditions on reaction rate and reaction mechanism. Meanwhile, validity of reaction mechanisms proposed for HDO of bio-oil model compounds could be confirmed by kinetic data.

In hydrodeoxygenation of phenol over Ni/HZSM-5 and Ni/Al2O3–HZSM-5 catalysts in aqueous media at 200 °C and 50 bar hydrogen pressure, reaction rates, turnover frequencies (TOF) and activation energies (Ea) of phenol hydrogenation to cyclohexanone, cyclohexanone hydrogenation to cyclohexanol, cyclohexanol dehydration to cyclohexene and cyclohexene hydrogenation to cyclohexane were determined.132 Rate of phenol hydrogenation was remarkably lower than that of cyclohexanone hydrogenation (Table 5). It was shown that Ni/Al2O3–HZSM-5 was five times more efficient than Ni/HZSM-5 for phenol hydrogenation. While, rate of cyclohexanone hydrogenation over Ni/Al2O3–HZSM-5 was only 1.5 times higher than that over Ni/HZSM-5 since Lewis acid sites of Al2O3 prevent from cyclohexanone hydrogenation and shift the equilibrium state of cyclohexanone ↔ cyclohexanol toward cyclohexanone. In aqueous reaction media, Brønsted acid sites are kinetically active in dehydration reaction while Lewis acid sites are almost inactive; higher amount of Brønsted acid sites in Ni/HZSM-5 (0.070 mmol g−1) compared to Ni/Al2O3–HZSM-5 (0.045 mmol g−1) resulted in higher rate of cyclohexanol dehydration. In another work for hydrogenation of phenol using mixed catalysts of Pd/C and H3PO4 liquid acid, it was revealed that dehydration of cyclohexanol was the rate determining step due to the low dehydration TOF of H3PO4 (15 mol molH+−1 h−1).27 Substitution of H3PO4 liquid acid with HZSM-5 solid acid enhanced dehydration TOF for 107 orders of magnitude.55 It was also shown that rate of cyclohexanol dehydration was considerably increased by impregnating Ni on HZSM-5; cyclohexanol dehydration to cyclohexene in an aqueous media is an equilibrium reaction and nickel addition to HZSM-5 promotes hydrogenation of cyclohexene and shifts the equilibrium reaction toward higher dehydration of cyclohexanol and formation of cyclohexene.132 Furthermore, temperature dependence of rate of phenol hydrogenation over Ni/HZSM-5 and Ni/Al2O3–HZSM-5 catalysts was investigated at reaction temperature range of 160–220 °C, and it was revealed that Ni/Al2O3–HZSM-5 was more dependent on temperature change compared to Ni/HZSM-5 due to its higher activation energy (56 kJ mol−1 on Ni/Al2O3–HZSM-5 and 48 kJ mol−1 on Ni/HZSM-5).

Table 5 Reaction rate, TOF and Ea data for aqueous phase phenol HDO reaction network at 200 °C over Ni/HZSM-5 and Ni/Al2O3–HZSM-5 catalystsa,132
Reaction Ni/HZSM-5 Ni/Al2O3–HZSM-5
a BAS: Brønsted acid sites.
Step 1: phenol hydrogenation
r1 (mmol g−1 h−1) 14 61
TOF1 (mol molsurf Ni−1 h−1) 398 553
Ea1 (kJ mol−1) 48 56
[thin space (1/6-em)]
Step 2: cyclohexanone hydrogenation
r2 (mmol g−1 h−1) 108 159
TOF2 (mol molsurf Ni−1 h−1) 2443 1233
Ea2 (kJ mol−1) 142 129
[thin space (1/6-em)]
Step 3: cyclohexanol dehydration
r3 (mmol g−1 h−1) 528 354
TOF3 (mol molBAS−1 h−1) 7428 8333
Ea3 (kJ mol−1) 112 114
[thin space (1/6-em)]
Step 4: cyclohexene hydrogenation
r4 (mmol g−1 h−1) 1813 2156
TOF4 (mol molsurf Ni−1 h−1) 55[thin space (1/6-em)]136 20[thin space (1/6-em)]287
Ea4 (kJ mol−1) 35 25

HDO of phenol over CoMoB amorphous catalyst in an autoclave reactor was described by pseudo-first-order reaction kinetics (eqn (8)).46

 
ln(1 − x) = kCcatt (8)
here x, k (mL g−1 s−1), t and Ccat refer to conversion of phenol, rate constant, reaction time, and concentration of catalyst in reactor at reaction time t, respectively. Temperature dependency of reactant consumption exhibited that increase of reaction temperature from 250 °C to 275 °C enhanced phenol conversion leading to increase in rate constant. Pseudo-first-order kinetic analysis was also used to study the effects of catalyst to phenol ratio, calcination temperature in catalyst preparation step, textural properties of catalyst and reaction temperature on phenol HDO rate over Pd supported on polymeric mesoporous graphitic carbon nitride (Pd/mpg-C3N4).126 Catalyst to phenol ratio remarkably affected reaction rate, and phenol conversion was enhanced from ∼9 to ∼50% by increase of catalyst to phenol ratio from 1.7 to 8.4. It was revealed that increase of calcination temperature (100–200 °C) reduced reaction rate constant due to the change of composition and crystalline structure of palladium species. Meanwhile, different values of BET surface area and mesopore volume of catalyst caused by different ratios of MSiO2/Mcyanamide (0.5–1.5) in mpg-C3N4 support resulted in remarkably different catalytic activities. However, product selectivity (100% cyclohexanone) was shown to be unchanged by textural change of catalyst. Furthermore, complete conversion of phenol was achieved at both reaction temperatures of 30 and 80 °C but at reaction times of 38 and 3 h, respectively, indicating the significant dependency of reaction rate on reaction temperature.

Sulman et al.145 studied catalytic performance of Pd supported on hypercrosslinked polystyrene (HPS) and Al2O3 in phenol hydrogenation, and proposed two possible reaction mechanisms for phenol transformation: (i) phenol hydrogenation to cyclohexanone (route 1) followed by cyclohexanone hydrogenation to cyclohexanol (route 2); (ii) direct formation of cyclohexanol through phenol hydrogenation (route 3). Each of these reaction routes has its own reaction rate. Assuming a constant hydrogen pressure during the reaction, kinetic equations for phenol conversion and cyclohexanone formation are as follows:

 
image file: c5ra22137d-t3.tif(9)
 
image file: c5ra22137d-t4.tif(10)
where x (mol%), τ (s), y (mol%), kΣ = k1 + k3 (cm3 g−1 s−1), k2 (cm3 g−1 s−1) and f(c) are phenol conversion, contact time, cyclohexanone yield, overall rate constant of phenol transformation through reaction routes 1 and 3, rate constant of reaction route 2 and function of concentrations of the reactants which determines the stationary concentrations of free active sites or intermediates on the catalyst surface. Dividing eqn (10) by (9) results in dimensionless eqn (11) which shows the relation between cyclohexanone yield and phenol conversion:
 
image file: c5ra22137d-t5.tif(11)
where s0 = k1/kΣ and b are initial selectivity at phenol conversion of zero and relative rate parameter of transformation of cyclohexanone to cyclohexanol, respectively. Integrating eqn (11) gives (12) which is used for selectivity (s) determination:
 
image file: c5ra22137d-t6.tif(12)

Meanwhile, ymax (mol%) and xmax (mol%) could be calculated through the expression eqn (13) in order to estimate catalyst efficiency.

 
image file: c5ra22137d-t7.tif(13)

In this study, cyclohexanone selectivity was shown to be a function of phenol conversion; selectivity toward cyclohexanone was significant at low phenol conversion over Pd/HPS and Pd/Al2O3, while it was reduced by increase of phenol conversion. Meanwhile, it was revealed that dependency of cyclohexanone selectivity on phenol conversion was not affected by change of reaction temperature using Pd/HPS since the activation energies for reaction routes 1–3 were almost similar over this catalyst. However over Pd/Al2O3, increase of reaction temperature led to higher reduction of cyclohexanone selectivity by increase of phenol conversion. In addition, increase of reaction temperature resulted in enhanced activity of Pd/HPS and lower phenol conversion over Pd/Al2O3 indicating that the activation energy for reactions differs over the two catalysts.

Mechanism of phenol alkylation over mixed catalysts of Pd/C with HBeta or La/HBeta zeolites in the presence of hydrogen was reported by Zhao et al.146 Cyclohexanol produced from phenol hydrogenation was involved in phenol alkylation over zeolite acid sites to form heavier products. Experimental data obtained from phenol hydroalkylation were fitted using kinetic equation such as:

 
image file: c5ra22137d-t8.tif(14)
where [A], [C], −rA (mol h−1), τ (h), kA (h−1) and kE (h−1) refer to phenol concentration, cyclohexanol concentration, rate of phenol consumption, reaction time, rate constant of phenol consumption and rate constant of formation of heavy compounds. It was shown in this research that the ratio of phenol to cyclohexanol controls rate of alkylation reaction; high ratio of phenol to cyclohexanol which is achieved due to the slow rate of phenol hydrogenation enhances the alkylation mechanism.

The effect of nickel particle size147 and nickel content98,128 on hydrogenation rate of phenol was investigated using Ni/SiO2 and Ni/HZSM-5 catalysts. Smaller average size of Ni particles resulted in lower specific rate of hydrogenation.147 Meanwhile, increase of nickel content in catalyst led to remarkable increase of reaction constant due to the change of catalyst structure; increase of Ni content resulted in increase of nickel particle size and surface area of exposed nickel.128 High content of Ni results in high accessibility to nickel active sites and enhances the adsorption and conversion of phenol on catalyst surface.98

The intrinsic activities of Ni2P/ZrO2, Ni2P/Al2O3 and Ni2P/SiO2 catalysts for HDO of guaiacol were compared at low conversion of guaiacol (<20%) using pseudo-first-order kinetic analysis.75 1.3, 1.5 and 1.2 L (g catalyst)−1 h−1 were the HDO reaction rates of Ni2P/Al2O3, Ni2P/ZrO2 and Ni2P/SiO2, respectively, indicating that highest activity was achieved using ZrO2 as support. Meanwhile, rate constant of guaiacol HDO over Pt/Al2O3 catalyst was 16.2 L (g catalyst)−1 h−1, illustrating that Pt is much more active than Ni2P for HDO of guaiacol.148 Bui et al.23 investigated the kinetics of HDO of guaiacol over unsupported and supported MoS2 and CoMoS catalysts at 300 °C in a fixed bed tubular reactor. Rate constant of reaction was analyzed based on a kinetic expression of the type:

 
image file: c5ra22137d-t9.tif(15)
here τ is the total conversion, w is the catalyst weight (g), Fo is the molar flow of reactant (mol s−1) and Co is the initial concentration of reactant (mol L−1). They also expressed the specific reaction rate r (mol g−1 s−1) and the intrinsic reaction rate ri (molec. Mo at−1 s−1) as the following equations:
 
image file: c5ra22137d-t10.tif(16)
 
ri = (r/n) × N (17)
where n is number of Mo atoms per g of catalyst and N is Avogadro number. Data analysis demonstrated that alumina supported CoMoS had higher total guaiacol conversion and lower HDO rate (Table 6). In fact, alumina support led to high conversion of guaiacol into catechol and methylated compounds instead of oxygen-free products.

Table 6 Conversion of guaiacol as well as specific and intrinsic HDO rates for unsupported and supported MoS2 and CoMoS catalysts at 300 °C[thin space (1/6-em)]23
Catalyst rguaiacol (10−7 mol g−1 s−1) rHDO (10−7 mol g−1 s−1) ri(guaiacol) (10−4 molec. Mo at−1 s−1) ri(HDO) (10−4 molec. Mo at−1 s−1)
MoS2 35 10.1 6 1.7
CoMoS 222 49.7 72 16.2
MoS2/Al2O3 64.2 2.3 81 2.9
CoMoS/Al2O3 71 5 93 6.6

On the other hand, a Langmuir–Hinshelwood model (eqn (18)–(20)) was used in the kinetic study of HDO of phenols over CoMoS/Al2O3 catalyst in a flow reactor at 300 °C and 28.5 bar hydrogen pressure.40,41

 
image file: c5ra22137d-t11.tif(18)
 
image file: c5ra22137d-t12.tif(19)
 
image file: c5ra22137d-t13.tif(20)
here A, B and C are mole fractions of phenolic compounds, benzene and hydrocarbons (cyclohexene and cyclohexane), respectively. KA is the equilibrium constant of adsorption of reactant A on catalyst, k1 is the rate constant of hydrogenolysis path to benzene production, k2 is the rate constant of hydrogenation path to hydrocarbon formation, Co is the concentration of feed A, τ is the space time variable and n is the inhibition order. The correlation between reactant concentration and space time indicated that HDO reaction was inhibited by the adsorption of reactant on catalyst.

The first-order kinetic model presented by Mortensen et al.149 for HDO of phenol revealed that although Ru/C, Pt/C and Pd/C are appropriate catalysts for hydrogenation, their performance in deoxygenation is poor (Table 7). They concluded that the best hydrogenation catalyst is not necessarily the best deoxygenation catalyst. Runnebaum et al.150 used pseudo-first-order kinetics to study the rate constant of HDO reactions of guaiacol and anisole (Table 7). They pointed out that different reactants have different reactivities; the rate constant of demethoxylation of guaiacol and anisole were 4.4 and 0.86 L (g of catalyst)−1 h−1, respectively. Indeed, the presence of hydroxy group in the ortho position of guaiacol molecule leads to easier removal of methoxy group. In addition, the rate constant of demethylation of methoxy substituent group of anisole (12 L (g of catalyst)−1 h−1) is higher than that of guaiacol (6.5 L (g of catalyst)−1 h−1). In another study, pseudo-first-order analysis was used for determination of rate constants of transformation of anisole to primary products of phenol, 2-methylphenol, 2-methylanisole and 4-methylphenol over Pt/Al2O3 catalyst.151 Rate constant of phenol production (12 L (g of catalyst)−1 h−1) was much higher than that of 2-methylphenol (2.8), 2-methylanisole (0.14) and 4-methylphenol (0.039). These data show that the primary reaction of anisole demethylation was more favorable than transalkylation of anisole. Very low rate of 4-methylphenol formation compared to formation rate of 2-methylphenol shows that Pt/Al2O3 kinetically proceeds transalkylation mechanism in ortho position.

Table 7 Pseudo-first-order rate constants for HDO of phenols
Reactant T (°C) P (bar) Catalyst Rate constant Reaction type Ref.
a Determined from a 18 min isothermal experiment at 275 °C with 50% phenol conversion.b Determined from a 10 min isothermal experiment at 275 °C with 99.8% phenol conversion.c Determined from a 16 min isothermal experiment at 275 °C with 44% phenol conversion.
Phenol 300 28.5 CoMoS/Al2O3 ∼8 mol (kg cat)−1 h−1 Hydrogenolysis 45
Phenol 300 28.5 CoMoS/Al2O3 ∼40 Hydrogenation  
m-Cresol 300 28.5 CoMoS/Al2O3 ∼20 Hydrogenolysis  
m-Cresol 300 28.5 CoMoS/Al2O3 ∼40 Hydrogenation  
p-Cresol 300 28.5 CoMoS/Al2O3 ∼32 Hydrogenolysis  
p-Cresol 300 28.5 CoMoS/Al2O3 ∼45 Hydrogenation  
Guaiacol 300 1.4 Pt/Al2O3 0.11 L (g cat)−1 h−1 Hydrodeoxygenation 150
Guaiacol 300 1.4 Pt/Al2O3 4.40 Hydrodeoxygenation  
Guaiacol 300 1.4 Pt/Al2O3 6.50 Hydrogenolysis  
Guaiacol 300 1.4 Pt/Al2O3 1.80 Transalkylation  
Guaiacol 300 1.4 Pt/Al2O3 0.50 Bimolecular transalkylation  
Guaiacol 300 1.4 Pt/Al2O3 0.21 Bimolecular transalkylation  
Guaiacol 300 1.4 Pt/Al2O3 0.26 Bimolecular transalkylation  
Anisole 300 1.4 Pt/Al2O3 12 Hydrogenolysis  
Anisole 300 1.4 Pt/Al2O3 0.86 Hydrodeoxygenation  
Anisole 300 1.4 Pt/Al2O3 2.8 Transalkylation  
Anisole 300 1.4 Pt/Al2O3 0.039 Transalkylation  
Anisole 300 1.4 Pt/Al2O3 0.14 Bimolecular transalkylation  
Phenol 275 100 Ru/C 1950a mL (kg cat)−1 min−1 Hydrogenation 149
Phenol 275 100 Pt/C 31[thin space (1/6-em)]200b Hydrogenation  
Phenol 275 100 Pd/C 1840c Hydrogenation  
Phenol 275 100 Ru/C 115 Deoxygenation  
Phenol 275 100 Pt/C 1 Deoxygenation  
Phenol 275 100 Pd/C 19 Deoxygenation  
Anisole 300   Pt/Al2O3 12.8 L (g cat)−1 h−1 Demethylation 108
Anisole 350   Pt/Al2O3 19.2 Demethylation  
Anisole 400   Pt/Al2O3 28.3 Demethylation  
Anisole 300   Pt/Al2O3 6.9 Transalkylation to 2-methylphenol  
Anisole 350   Pt/Al2O3 12.8 Transalkylation to 2-methylphenol  
Anisole 400   Pt/Al2O3 39.6 Transalkylation to 2-methylphenol  
Anisole 300   Pt/Al2O3 5.8 Transalkylation to 2,6-dimethylphenol  
Anisole 350   Pt/Al2O3 13.7 Transalkylation to 2,6-dimethylphenol  
Anisole 400   Pt/Al2O3 22.3 Transalkylation to 2,6-dimethylphenol  
Anisole 300   Pt/Al2O3 2.4 Transalkylation to 2,4,6-trimethylphenol  
Anisole 350   Pt/Al2O3 3.5 Transalkylation to 2,4,6-trimethylphenol  
Anisole 400   Pt/Al2O3 8.5 Transalkylation to 2,4,6-trimethylphenol  

6. Transfer hydrogenation of phenol

Severe operating conditions of HDO of bio-oil to high value chemicals make it difficult to industrialize the process. Flammability of molecular hydrogen in contact with air leads to hard control of high pressure of hydrogen in large scale industrial applications of HDO process. Besides, high price of hydrogen gas production and compression as well as the difficulties related to its transportation and storage limit the application of hydrogen gas.152,153 Therefore, replacing hydrogen gas with liquid hydrogen donors (H-donor) as alternative source of hydrogen is highly attractive. Hydrogen released from H-donor compounds could be reacted with hydrogen acceptors on catalytic active sites. Catalytic reduction of a molecule using hydrogen atom transferred from donor molecule is called transfer hydrogenation.154–156 Transfer hydrogenation is occurred at less severe operating condition compared to hydrogenation using hydrogen gas. The precise mechanism of catalytic transfer hydrogenation is not yet clear. Generally, hydrogenation could occur in two mechanisms: (i) decomposition of H-donor to hydrogen molecule which is activated on catalytic active sites for involving in hydrogenation reaction; (ii) transfer of hydrogen atom from H-donor molecule to acceptor on catalytic sites. The general reaction between donor (DHx) and acceptor (A) is supposed to be like eqn (21).157
 
image file: c5ra22137d-t14.tif(21)

The compound used as hydrogen donor should have high reduction potential through which hydrogen transfer could be occurred in mild operating conditions. Meanwhile, compounds formed via hydrogen release of H-donors should be easily separated from final products obtained from transfer hydrogenation process. So far, hydroaromatics,158,159 unsaturated hydrocarbons,160,161 alcohols,162–165 acids166–168 and formate salts155,169,170 have been used as hydrogen donor for transfer hydrogenation reactions. Among all types of donors, alcohols are most attractive since they are non-corrosive and act as both reactants and solvents, and also their dehydrogenation products (aldehydes and ketones) are easily separated from final products.152 Secondary alcohols such as 2-propanol, 2-butanol, 2-pentanol, 2-heptanol and 2-octanol are more efficient than primary alcohols such as ethanol, 1-propanol and 1-butanol in transfer hydrogenation reactions.156 Methanol, ethanol, 2-propanol and 2-butanol were examined in transfer hydrogenation of phenol over RANEY® Ni;154 complete conversion of phenol was only achieved by the use of 2-propanol and 2-butanol, while methanol and ethanol led to catalyst deactivation (see Table 8, entries 1–8). Contradictory data achieved using different types of alcohols were assigned to different interactions between catalyst and donors. Adsorption of methanol and ethanol on Ni surface splits the O–H bond producing an H-adatom and surface alkoxy groups which block the surface of RANEY® Ni and deactivate the catalyst. On the other hand, 2-propanol adsorption on Ni surface cleaves the both O–H and α-C–H bonds releasing two atoms of H which are adsorbed on the catalyst surface and acetone is produced instead of alkoxy groups. These reactions lead to the loading of catalyst surface with H-atoms and keep the catalyst surface unblocked for transfer hydrogenation. In addition to alcohols, formic acid, a major product of biomass processing, has high potential to be used as hydrogen donor in hydrogenation.171 Formic acid is an inexpensive, environmentally non-malignant and easy to handle hydrogen source.172 Formate salts are also considered as important H-donors since they are highly capable for hydrogen production, stable and easily accessible. Furthermore, bicarbonate which is produced from dehydrogenation of formate could be converted back to formate at mild conditions.155 In transfer hydrogenation of phenol over Pd/C catalyst, hydrogen donor of sodium format/water was reported to result in the selectivities of 98.1 and 1.9% for cyclohexanone and cyclohexanol, respectively, at 63.6% phenol conversion.173 Potassium formate was also used in transfer hydrogenation of phenol over Pd/C catalyst which resulted in 85% phenol conversion and ∼100% cyclohexanone selectivity.153

Table 8 Transfer hydrogenation of phenol at different conditions
Entry H-donor Catalyst T (°C) Time (h) Conv. (%) Selectivity (%) Ref.
Cyclohexanone Cyclohexanol
1 2-Propanol RANEY® Ni 80 3 100 0.4 99.3 154
2 2-Propanol RANEY® Ni 120 3 100 0.7 98.7  
3 2-Butanol RANEY® Ni 80 3 100 1.1 98.2  
4 2-Butanol RANEY® Ni 120 3 100 0.8 98.8  
5 Methanol RANEY® Ni 80 3 18 5.3 93.5  
6 Methanol RANEY® Ni 120 3 21 4.8 94.1  
7 Ethanol RANEY® Ni 80 3 15 7.1 92.9  
8 Ethanol RANEY® Ni 120 3 20 5.1 91.7  
9 HCOONa Pd/C 80 2 32.2 99.8 0.2 173
10 HCOONa Pd/C 80 5 63.6 98.1 1.9  
11 Formic acid Pd/C (commercial) 50 4 56 98.6 1.4 171
12 Formic acid Pd/AC 50 4 65.6 96.3 3.7  
13 Formic acid Pd/MIL-101 50 4 37.9 98.2 1.8  
14 Formic acid Pd/TiO2 50 4 24.7 99.2 0.8  
15 Formic acid Pd/Al2O3 50 4 18.1 99.4 0.6  
16 Formic acid Pd/TiO2–AC 50 4 47.3 96.4 3.6  
17 2-Methyltetrahydrofuran RANEY® Ni 80 3 15 7.9 91.2  
18 Methanol Ni/SiO2 0 5 min 20 — 100 180
19 HCOOK Pd/C 100 5 85 >99 0 153
20 HCOOK Pd/C 100 6 >99 >99 0  
21 HCOOK Pd/C 100 7 >99 95 5  
22 HCOOK Pd/C 100 14 >99 75 25  
23 HCOOK Pd/C 70 6 94 97 3  
24 HCOOK Pd/C 50 6 19 >99 0  

In addition to H-donors, catalyst characteristics have also remarkable effect on deoxygenation efficiency of transfer hydrogenation process. So far, homogeneous catalysts have been vastly used in transfer hydrogenation reactions.174–176 Difficult reusability of homogeneous catalysts and the necessity for a base or ligand in homogeneous catalysis make it difficult to handle these catalytic systems. Heterogeneous catalysts do not have such problems and are more appropriate for industrial application. Heterogeneous catalysts of Pd/C, Pd/Al2O3, Pd–La/Al2O3, Ru/C, RANEY® Ni and Ni–Cu/Al2O3 were used in transfer hydrogenation of phenols, aldehydes, ketones and bio-oil.153,155,177–179 It is generally supposed that both the reactions of hydrogen liberation from H-donor and hydrogen consumption by acceptor are simultaneously occurred on the same catalytic active sites. Therefore, an efficient catalyst for transfer hydrogenation reaction should have high potential for adsorption of both hydrogen donor and acceptor molecules. For instance, in transfer hydrogenation of phenol using formic acid as H-donor, Pd supported on AC had higher activity compared to Pd supported on MIL-101, TiO2, Al2O3 and TiO2–AC due to the facilitating effect of activated carbon on adsorption of both phenol and formic acid (see Table 8, entries 12–16).171 At 65.6% phenol conversion over Pd/AC, the use of formic acid as hydrogen donor led to production of cyclohexanone and cyclohexanol with selectivities of 96.3 and 3.7%, respectively.171 Meanwhile, the rates of hydrogen release from H-donor and hydrogen consumption by acceptor should be synergistic since the hydrogenation of acceptor molecule is occurred by the hydrogen liberated from H-donor molecule. Higher rate of hydrogen liberation compared to hydrogen consumption results in inefficient consumption of released hydrogen while a fraction of generated hydrogen is not involved in hydrogenation reaction. On the other hand, too slow rate of hydrogen release by H-donor in comparison with hydrogen consumption by acceptor could lead to reduction in hydrogenation of acceptor molecule due to the polymerization of acceptor prior to its hydrogenation on catalytic active sites. Further understanding of the interaction between catalyst and H-donor is essential for development of transfer hydrogenation of bio-oil/bio-oil model compounds.

7. Conclusions and recommendations for future work

Several researches have been done in order to explore efficient catalytic systems for selective hydrodeoxygenation of bio-oil phenolic compounds to high-value hydrocarbons. Phenols are highly stable in HDO reaction and their oxygen removal is difficult due to the strong bond between Caromatic and O. HDO reaction could be proceeded through different pathways leading to the formation of different products. Selective hydrodeoxygenation of phenols to valuable hydrocarbons is a strong function of catalyst properties and operating conditions. The effects of promoters, support materials, catalyst preparation method, operating conditions, solvent and co-feeding on catalytic performance and reaction selectivity of HDO of phenol, cresol, guaiacol and anisole have been reviewed in this study. Generally, multifunctional catalysts are more suitable than monofunctional catalysts in HDO of phenols. Metals applied in HDO process should have high potential for activating both oxygenates and hydrogen molecules. Selection of metals which are in good interaction with each other can improve the dispersion of active sites in catalyst and prevent their agglomeration. From this review, it is inferred that metal promoters like Mo, Ni, Co, La and Cu as well as non-metal promoter of phosphorus increase the selectivity toward oxygen-free products in HDO of phenol, guaiacol and anisole. Use of an appropriate support could enhance the interaction between reactants and catalyst surface and reduce coke formation. A support with high surface area like activated carbon can cause more active functional groups on catalyst surface and better dispersion of metal active sites leading to enhanced catalytic activity. Although acidic supports such as alumina and zeolites can improve HDO rate, but the high coke formation caused by these supports reduces catalyst lifetime. Method of catalyst preparation is influential on catalytic performance and reaction selectivity by affecting catalyst properties such as active metal dispersion, particle size and morphology. Since hydrodeoxygenation and hydrogen adsorption are exothermic reactions, reaction temperature is needed to be optimized in order not to restrict these reactions. Besides, high hydrogen pressure is required to improve the efficiency of HDO process. The solvent selected for HDO of phenols should improve the miscibility of both oxygenates and hydrogen and eliminate mass transfer limitations in order to increase diffusion rate of reactants in catalyst. Considering the fact that HDO could be proceeded through a variety of reaction pathways which result in formation of different products, further researches are needed in order to develop efficient catalysts and to optimize reaction conditions for selective hydrodeoxygenation of phenolic compounds to high-value hydrocarbons. The following recommendations are suggested for further research on catalytic transformation of phenolic model compounds of bio-oil to high-value chemicals through hydrogenation:

• Exploration of functionality of different metals for selective conversion of phenolic compounds into hydrocarbons of aromatics and alicyclics through hydrogenolysis and hydrogenation reactions, respectively.

• Optimization of ratio of metal to acid sites in bifunctional metal/acid catalysts.

• Use of pore- and acid-modified activated carbon as catalyst support; supports with mild acidity and mesopore structure could improve HDO efficiency.

• Study of the effect of solvent type on synergistic rates of hydrogen release from H-donor and hydrogen consumption by acceptor in transfer hydrogenation of phenolic model compounds of bio-oil.

Acknowledgements

The authors wish to acknowledge the Faculty of Engineering at University of Malaya for financial support through the HIR Grant (D000011-16001).

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