Production and isolation of monocyclic aromatic compounds from a gasification gas

Abstract

The present invention concerns a process and system for producing and isolating a fraction of monocyclic aromatic compounds from a gasification gas. The process comprises (a) contacting the gas with a catalyst capable of converting ethylene and possibly other unsaturated hydrocarbons into monocyclic aromatic compounds; and (b) isolating monocyclic aromatic compounds from the gas originating from step (a). The present invention is ideally suited for treatment of gas from coal, biomass or waste gasification, which comprises substantial amounts of ethylene as well as monocyclic aromatic compounds. Treatment according to the invention first converts the ethylene into further monocyclic aromatic compounds, and the entire fraction of monocyclic aromatic compounds is isolated to obtain a valuable product.

Claims

1. A process for producing a fraction of monocyclic aromatic compounds from a product gas of a gasification process, comprising: (a) contacting the gas with a catalyst capable of converting ethylene into monocyclic aromatic compounds; (b) isolating monocyclic aromatic compounds from the gas originating from step (a).

2. The process according to claim 1, wherein the gasification process involves the gasification of coal, biomass or waste.

3. The process according to claim 1, wherein the isolating of step (b) is performed by (b1) contacting the gas originating from step (a) with a washing liquid, at a temperature of 15-60° C., to obtain the purified gas and a spent washing liquid; (b2) contacting the spent washing liquid with a stripping gas comprising steam, to obtain a loaded stripping gas comprising monocyclic aromatic compounds and a stripped washing liquid; and (b3) separating the monocyclic aromatic compounds from the loaded stripping gas obtained in step (b2) to obtain a composition comprising monocyclic aromatic compounds.

4. The process according to claim 3, wherein the stripping gas comprises at least 95 vol. % steam.

5. The process according to claim 3, wherein the process further comprises step (c) that involves condensation of the loaded stripping gas to obtain an immiscible liquid composition comprising water and monocyclic aromatic compounds, collection of the liquid composition in a vessel and liquid-liquid separation thereof.

6. The process according to claim 3, wherein the washing liquid comprises an organic polysiloxane.

7. The process according to claim 1, wherein the monocyclic aromatic compounds include one or more selected from benzene, toluene, xylenes and ethylbenzene.

8. The process according to claim 1, wherein the catalyst capable of converting ethylene into monocyclic aromatic compounds comprises a zeolite, wherein the zeolite is optionally a ZSM-5 zeolite, and wherein the zeolite is optionally promoted with Ga, Zn and/or Mo.

9. The process according to claim 1 any one of the preceding claims, wherein step (a) is performed at a temperature of 250-650° C.

10. The process according to claim 1, wherein the gas that is subjected to step (a) comprises ethylene and optionally at least one of (i) 5-30 vol % CH.sub.4; (ii) 1-15 vol % C.sub.2H.sub.x, wherein x=2, 4 or 6; (iii) 1-10 vol. % C.sub.yH.sub.z, wherein y=3, 4 or 5 and z=(2y−2), (2y) or (2y+2); (iv) 10-60 vol % H.sub.2; (v) 5-50 vol % CO; and (vi) 5-50 vol % CO.sub.2, based on total dry volume.

11. The process according to claim 1, wherein tar-like components are removed from the gas prior to step (a), by: (c1) contacting the gas with a pre-washing liquid at a temperature of 60-150° C., to obtain a detarred gas which is fed to step (a) and a spent pre-washing liquid; and (c2) contacting the spent pre-washing liquid with a tar stripping gas, to obtain a loaded tar stripping gas and a stripped pre-washing liquid.

12. The process according to claim 1, wherein the gas is subjected to water removal prior to step (a), wherein the water removal is optionally conducted in a condenser.

13. A modular system for performing the process according to claim 1, comprising: (a) an ethylene conversion module for converting ethylene into monocyclic aromatic compounds, comprising a gas inlet (a1) for receiving the gas, a catalyst (a2) capable of converting ethylene into monocyclic aromatic compounds and a gas outlet (a3) for discharging a gas enriched in monocyclic aromatic compounds; (b1) an absorbing unit comprising a gas inlet (b11) for receiving the gas enriched in monocyclic aromatic compounds, a liquid inlet (b12) for receiving a washing liquid, a gas outlet (b14) for discharging a purified gas and a liquid outlet (b15) for discharging a spent washing liquid; and (b2) a stripping unit, comprising a liquid inlet (b21) for receiving the spent washing liquid, a gas inlet (b22) for receiving a stripping gas, a gas outlet (b23) for discharging a loaded stripping gas and a liquid outlet (b24) for discharging a stripped washing liquid, wherein outlet (a3) is in fluid connection with inlet (b11), outlet (b15) is in fluid connection with inlet (b21) and wherein outlet (b24) is optionally in fluid connection with inlet (b12).

14. The modular system according to claim 13, further comprising: (b3) a separating module, comprising a gas inlet (b31) for receiving the loaded stripping gas, means (b32) for separating the monocyclic aromatic compounds from the stripping gas, an outlet (b33) for discharging a cleared stripping gas, and an outlet (b34) for discharging the monocyclic aromatic compounds, wherein outlet (b23) is in fluid connection with inlet (b31) and wherein outlet (b33) is optionally in fluid connection with inlet (b22).

15. The modular system according to claim 13, further comprising: (c1) a pre-washing unit comprising a gas inlet (c11) for receiving a gas comprising tar-like components and monocyclic aromatic compounds, a liquid inlet (c12) for receiving a pre-washing liquid, a gas outlet (c14) for discharging a detarred gas and a liquid outlet (c15) for discharging a spent pre-washing liquid; and (c2) a tar stripping unit, comprising a liquid inlet (c21) for receiving the spent pre-washing liquid, a gas inlet (c22) for receiving a tar stripping gas, a gas outlet (c23) for discharging loaded tar stripping gas and a liquid outlet (c24) for discharging a stripped pre-washing liquid, wherein outlet (c14) is in fluid connection with inlet (al), outlet (ci5) is in fluid connection with inlet (c21) and wherein outlet (c24) is optionally in fluid connection with inlet (c12).

Description

DESCRIPTION OF THE FIGURES

(1) FIG. 1A depicts a preferred embodiment of the process and system according to the invention, with reference to the description of the system above and accompanying reference numbers. FIG. 1B depicts a preferred embodiment of separation module (b3), with reference to the description of the system above and accompanying reference numbers.

(2) FIG. 2 depicts the configuration inside the reactor used in Example 1. (A) alumina beads; (Q) quartz wool; (C) catalyst material; T=temperature indicators; (1) feed gas; (2) N.sub.2/H.sub.2/air inlet; (3) outlet gas; (4) condensate; (5) remaining gas to the afterburner; (6) dry gas to gas analysis.

(3) FIGS. 3-5 depict the effects of temperature on inlet and outlet concentrations (in vol %, based on dry volume), as obtained in Example 1. (I) inlet gas; (A) reactor at 500° C.; (B) reactor at 400° C.; (C) reactor at 300° C.; (A*) reactor at 500° C. with NH.sub.3 present in the feed. FIG. 3 shows the concentrations [C] of CO, H.sub.2, CO.sub.2 and CH.sub.4, FIG. 4 shows the concentrations of C.sub.2H.sub.x (x=2, 4 or 6) and FIG. 5 of C.sub.2H.sub.4 [E], benzene [B] and toluene [T].

(4) FIG. 6 depicts the effects of temperature on ethylene conversion (EC) and carbon selectivity (CS) to benzene (B) and toluene (T), as obtained in Example 1. (A) reactor at 500° C.; (B) reactor at 400° C.; (C) reactor at 300° C.

(5) FIG. 7 shows the effect of temperature on the concentration [C] in mg/Nm.sup.3 dry gas of aromatic compounds other than benzene and toluene in the inlet gas (I) and outlet gas (T=300° C.; T=500° C.) measured by SPA analysis, as obtained in Example 1. (1) ethylbenzene; (2) m/p-xylene; (3) o-xylene+styrene; (4) phenol; (5) indene+o-cresol; (6) m/p-cresol; (7) naphthalene; (8) quinolone; (9) isoquinoline; (10) 2-methylnaphthalene; (11) 1-methylnaphthalene; (12) biphenyl; (13) 2-ethylnaphthalene; (14) acenaphtylene; (15) acenaphtene; (16) fluorine; (17) phenanthene; (18) anthracene; (19) pyrene; (20) benzo(a)anthracene; (21) chrysene; (22) benzo(b)fluoranthene; (23) benzo(k)fluoranthene; (24) benzo(e)pyrene; (25) benzo(a)pyrene; (26) indeno(123-cd)pyrene; (27) dibenz(ah)anthracene; (28) benzo(ghi)perylene; (29) coronene.

(6) FIGS. 8-10 depict the effects of temperature on inlet and outlet concentrations (in vol %, based on dry volume), as obtained with 2.5% Ga ZSM-5 catalyst in Example 1. (I) inlet gas; (R) reactor in operation. FIG. 8 depicts the inlet and outlet concentrations of ethylene [E], benzene [B] and toluene [T]. FIG. 9 depicts the concentrations of C.sub.2H.sub.x (x=2, 4 or 6) and FIG. 10 the concentrations [C] of CO, H.sub.2, CO.sub.2 and CH.sub.4.

(7) FIG. 11 depicts the effect of the Ga loading on ethylene conversion (EC), as obtained in Example 1.

(8) FIGS. 12 and 13 depict the effect of the Ga loading on carbon selectivities (CS), as obtained in Example 1. In FIG. 12, the selectivities to benzene (FIG. 12A) and toluene (FIG. 12B) are shown, and in FIG. 13 the selectivities to benzene+toluene (FIG. 13A) and ethane (FIG. 13B).

(9) FIG. 14 shows the effect of Ga loading on the concentration [C] in mg/Nm.sup.3 dry gas of aromatic compounds other than benzene and toluene in the inlet gas (I) and outlet gas (0 wt %, 0.5 wt % and 2.5 wt % Ga loading) measured by SPA analysis, as obtained in Example 1. (1) ethylbenzene; (2) m/p-xylene; (3) o-xylene+styrene; (4) phenol; (5) indene+o-cresol; (6) m/p-cresol; (7) naphthalene; (8) quinolone; (9) isoquinoline; (10) 2-methylnaphthalene; (11) 1-methylnaphthalene; (12) biphenyl; (13) 2-ethylnaphthalene; (14) acenaphtylene; (15) acenaphtene; (16) fluorine; (17) phenanthene; (18) anthracene; (19) pyrene; (20) benzo(a)anthracene; (21) chrysene; (22) benzo(b)fluoranthene; (23) benzo(k)fluoranthene; (24) benzo(e)pyrene; (25) benzo(a)pyrene; (26) indeno(123-cd)pyrene; (27) dibenz(ah)anthracene; (28) benzo(ghi)perylene; (29) coronene.

(10) FIG. 15 depicts the carbon selectivities for toluene (I), benzene (II), m+p-xylene (III), o-xylene (IV) and ethyl-benzene (V), as obtained in Example 4.

EXAMPLES

(11) The examples below demonstrate the invention. In Example 1, ethylene present in a product gas is successfully converted into BTX. In Example 2, a BTX fraction is successfully isolated from a product gas. Example 3 describes the performance of aromatic harvesting by the BTX scrubbing unit. Example 4 illustrates the performance of the aromatization catalyst in a duration test.

Example 1

Experimental

(12) The GaZSM-5 catalysts were prepared by slurry wet impregnation of NH.sub.4—ZSM-5 (SiO.sub.2/Al.sub.2O.sub.3 molar ratio=30, CBV 3024E Zeolyst International) with aqueous solutions containing the appropriate amount of Ga(NO.sub.3).sub.3 (Alfa Aesar, 99.9%). Catalyst loadings of ca. 0.5 and 2.5 wt % Ga were prepared. The resulting materials were vacuum dried (70 mbar) for overnight at 60° C. and further calcined at 550° C. for 5 h. All the catalysts were pelletized and sieved to 40/70 mesh before testing.

(13) The reactor loaded with the catalyst was heated to 500° C. at a heating at a rate of 2° C./min under 0.5 L/min of a gas mixture composed of 60 vol. % H.sub.2 in N.sub.2. The H.sub.2/N.sub.2 activation gas was applied overnight. After intermediate N.sub.2 flushing, 0.5 L/min air was applied for 0.5 hours. After flushing again the reactor with N.sub.2, product gas from a MILENA gasifier, cleaned by an OLGA tar removal system and a gas cooler, was fed to the reactor. The 25 kWth MILENA gasifier (see: C. M. van der Meijden, Development of the MILENA gasification technology for the production of Bio-SNG. PhD. Thesis, 2010) was operated under the following operating conditions: ˜5 kg/h beech wood as biomass fuel, olivine as bed material, ˜850° C. gasification temperature, 1000 g/h steam fluidization, and 10 NmL/min neon injected as tracer gas in the settling chamber of the gasifier. The gasification system operated at atmospheric pressure. A slipstream of about 1 Nm.sup.3/h dry gas from MILENA was directed to the system downstream. OLGA (see: Dahlman Renewable Technologies, OLGA technology (2013); http://www.royaldahlman.com/renewable/home/tar-removal/olga-technology). After OLGA tar removal, most of the water contained in the gas was removed in a gas cooler operating at 5° C. Although about 90% of the ammonia contained in the gas is removed in the condensed water, the remaining traces of ammonia in the feed gas were further removed in a flask containing a 1 M nitric acid solution.

(14) The Ga-zeolite catalyst materials were tested in terms of activity and stability under relevant gasification conditions. The experiments were carried out in an oven fixed-bed reactor (28 mm diameter, 600 mm height) surrounded by an electrical oven. In all tests, both the height of the catalyst bed was set at 6 cm (20 g of catalyst), and the gas flow to 0.5 L/min (at atmospheric pressure) to keep the gas velocity similar. The configuration of the catalyst inside the reactor is plotted in FIG. 2. Above the catalyst bed (C), alumina beads (A) are placed for feed gas preheating. Four thermocouples (T) measure the temperature profile within the catalyst bed. Pressure indicators located before and after the reactor track the pressure drop over the bed. The gas composition at the inlet and outlet of the reactor was online measured using micro-GC analysis (Varian CP4900, with 3 columns with corresponding TCD detectors).

(15) From the molar balances performed over the reactor, several parameters have been calculated in order to assess the performance of the catalysts. Firstly, ethylene conversion is defined as:

(16) $\mathrm{Ethylene}\mathrm{conversion}\left(\%\right)=\frac{{\stackrel{.}{n}}_{C_2{H}_4,in}-{\stackrel{.}{n}}_{C_2{H}_4,\mathrm{out}}}{{\stackrel{.}{n}}_{C_2{H}_4,in}}\times 100$
where {dot over (n)} is the molar flow in mol/h. The inlet value has been taken in all cases as the last micro-GC analysis before switching to the outlet gas composition. On the other hand, carbon selectivity, i.e. the increase in the content of carbon contained in the product compound of generic formula C.sub.xH.sub.y with respect to the total amount of carbon converted from ethylene and acetylene, has been evaluated according to:

(17) $\mathrm{Carbon}\mathrm{selectivity}\mathrm{to}\mathrm{product}{C}_x{H}_y\left(\%\right)=\frac{x\left({\stackrel{.}{n}}_{C_x{H}_y,\mathrm{out}}-{\stackrel{.}{n}}_{C_x{H}_y,in}\right)}{2\left({\stackrel{.}{n}}_{C_2{H}_4,in}-{\stackrel{.}{n}}_{C_2{H}_4,\mathrm{out}}\right)+2\left({\stackrel{.}{n}}_{C_2{H}_4,in}-{\stackrel{.}{n}}_{C_2{H}_4,\mathrm{out}}\right)}\times 100$
where n.sub.i represents the molar flow of compound i in the gas (in mol/h), and x is the number of moles of carbon contained in the generic compound with formula C.sub.xH.sub.y.
Results

(18) The catalyst containing 2.5 wt. % Ga was analysed at a temperature of 300° C.-500° C. FIG. 3 presents the concentration of the main compounds in the product gas (CO, CO.sub.2, H.sub.2, and CH.sub.4). An initial shift was observed toward higher CO, and lower CO.sub.2 and H.sub.2 concentrations. Since the water content of the feed gas is very low—most of the water has been previously removed in the gas cooler—the water gas shift reaction is shifted toward the formation of CO and H.sub.2O. Lower temperatures favour the rate of the direct water-gas shift reaction. Towards the end of the test, a final stage wherein NH.sub.3 is present in the feed was briefly tested at 500° C. No hints of deactivation were detected. FIG. 4 displays the results of the C.sub.2H.sub.x concentration in the inlet and outlet gases. Complete conversion of acetylene was achieved, regardless of the reaction temperature. This is consistent with its highest reactivity among C.sub.2H.sub.x compounds. Moreover, lower temperatures lead to higher ethylene conversion, but lower ethane production. Ethane and methane might be by-products of the hydrogen-transfer mechanisms, in which the formation of diene, cyclic diolefins and aromatics is balanced by the formation of alkanes. The concentration of benzene and toluene in the outlet gas as well as the concentration of ethylene at various reaction temperatures, compared with the inlet gas concentrations, is depicted in FIG. 5. Benzene and toluene show distinct trends. A lower reactor temperature clearly reduces the benzene concentration in the outlet gas, showing that benzene production is favoured at higher temperatures. However, the decrease of toluene concentration is only evident when decreasing the temperature from 400° C. to 300° C. Although lower temperatures lead to lower benzene and toluene production, higher ethylene conversions were observed. Thus, a trade-off occurs between the conversion and selectivity to aromatics. No changes were observed in the gas composition when ammonia was present in the feed gas. The activation energy is highest for the final step of separation between the Ga.sup.+ active site and the benzene product, such that the separation process requires high temperature to enhance the reaction rate, which is in consistent with the obtained results. Furthermore, these results indicate different reaction mechanisms and active sites of toluene and benzene. Indeed, toluene formation is not directly related to benzene. In other words, the formation mechanism of toluene does not need benzene as an intermediate, suggesting that their active sites are not at the same location.

(19) FIG. 6 summarizes the effects of the reactor temperature on the conversion and carbon selectivity of the 2.5 wt. % Ga-ZSM-5 catalyst. Although decreasing the temperature from 500° C. to 300° C. slightly increases the ethylene conversion from 95% to 97%, it also results in lower carbon selectivity to aromatics. The carbon selectivity to benzene decreases from about 30% to even negative values when the temperature was decreased from 500° C. to 300° C., whereas the lowering of selectivity to toluene becomes evident when the temperature was decreased from 400° C. to 300° C., with values halving from over 30% to about 15%. These results mean that the composition of the fraction of monocyclic aromatic compounds can be fine-tuned by varying the conditions employed in step (a). Further, it is noteworthy observing that the NH.sub.3 content in the feed gas (tested during the last 500° C. stage) did not seem to affect the catalyst performance.

(20) Complementary to online micro-GC analysis, solid phase adsorption (SPA) analyses were also performed at the inlet and outlet gases under stable conditions for the determination of the content and composition of aromatic compounds which, unlike benzene and toluene, cannot be measured online through micro-GC. However, owing to the high volatility of benzene and toluene in the adsorption cartridge, the SPA quantification of benzene and toluene is not reliable, thus benzene and toluene are not reported in the SPA results. The micro-GC results are used instead for calculations. FIG. 7 shows the SPA results of the inlet gas and the outlet gases at 300° C. and 500° C. The catalytic reaction produces besides benzene and toluene other aromatic compounds such as ethylbenzene, xylenes, indene, naphthalene and methyl naphthalene. Due to the fact that o-xylene and styrene have similar retention times, mass spectrometry (MS) was complementary used to determine the contribution of each species to the peak area. MS analysis revealed that about 80% of the signal detected in the inlet syngas corresponded to styrene (i.e. 20% to o-xylene), whereas about 80% of the signal in the outlet gas corresponds to o-xylene. As can be observed, the reaction temperature influences dramatically the distribution of the aromatic products. Whereas lower temperatures favour the production of ethylbenzene and xylenes, higher temperatures give a slightly higher production of naphthalene and naphthalene derivatives.

(21) FIGS. 8-10 depict the steady-state situation of a 4-hour run at 500° C. with the 2.5 wt. % Ga-ZSM-5 catalyst. Similar stable conditions were achieved with the other catalyst loadings tested. As can be seen in FIG. 8, the ethylene concentration dropped from 3.5 vol. % dry to <0.1 vol. %, resulting in 97% conversion. The behaviour of the catalyst material was highly stable over the duration of the experiment. The feed product gas composition might lead to milder reducing conditions than, for instance, pure hydrocarbon feeds. Strong reducing atmospheres were reported as a plausible reason for the increased deactivation rate of Ga catalysts. The benzene concentration increased from about 7000 ppmv to over 10000 ppmv, whereas the toluene concentration increased from about 500 ppmv to 4000 ppmv (all concentrations expressed in dry basis). The temperature near the catalyst bed surface increased by about 13° C. upon starting the operation. FIG. 9 shows the complete conversion of acetylene, as well as a marked increase in the ethane concentration from about 0.25 vol. % to more than 1 vol. %, which increased overtime. The concentrations of the major product gas compounds (CO, H.sub.2, CO.sub.2, and methane), as depicted in FIG. 10, show a shift in the composition upon the start of the reactor operation toward higher CO concentrations and lower H.sub.2 and CO.sub.2 concentrations, which indicates some transient reverse WGS activity. However, over time H.sub.2 and CO.sub.2 progressively increase, and CO decreases until recovering similar concentrations to those in the inlet gas. The presence of CO and CO.sub.2 in the feed gas may contribute in a positive way to the catalyst stability. This hypothesis is consistent with the mild carbon formation (greyish catalyst) observed during post-inspection. The CH.sub.4 concentration slightly increases from 11 vol. % dry to about 12-13 vol. %.

(22) FIG. 11 shows the effect of loading amount of Ga on the conversion efficiency of ethylene. Higher Ga loadings favour ethylene conversion. The reference unloaded zeolite yields ethylene conversion of about 80%. As a comparison, Qiu et al., in Catal. Letters 52 (1998) 37-42, reported an ethylene conversion of only 40% at 520° C. with unloaded H-ZSM-5 zeolite by using 3 vol. % ethylene in methane as the feed gas. The highest ethylene conversion value, obtained with 2.5 wt. % Ga-zeolite, was 97%, which is slightly higher than previously reported values by Qiu et al. (93% conversion with 5 wt. % Ga loading). At a Ga loading of 0.5 wt. %, the ethylene conversion can be maintained as high as 85%-90%. Moreover, the 2.5 wt. % Ga— catalyst exhibits a considerably high stability for 4 hours. FIG. 12 shows the results of the effect of Ga loading on the carbon selectivity to benzene and toluene. The carbon selectivity to benzene markedly increases from 8% (zeolite without added Ga) to 32% (0.5 wt. % Ga). In all cases, the catalyst activity remained fairly stable over time. The 0.5 wt. % Ga-catalyst yields the highest selectivity to benzene. Moreover, the zeolite with Ga loading takes approximately 1 hour in stream until reaching stable conditions, whereas the addition of 0.5 wt. % Ga to the zeolite reduces the stabilization time to about 0.5 hour. Therefore, the loaded Ga seems also to improve the catalyst stability. The unloaded ZSM-5 sample exhibits the lowest carbon selectivity to benzene (about 8%), but the highest carbon selectivity to toluene (about 45%). Both the 0.5 wt. % and the 2.5 wt. % Ga catalysts show similar carbon selectivity to toluene (about 30%). FIG. 13A shows the overall carbon selectivity to toluene and benzene. Although the addition of Ga leads to reduced carbon selectivity to toluene, the overall carbon selectivity to toluene and benzene still increases. Ethane formation is favoured at higher Ga contents in the catalyst (see FIG. 13B).

(23) FIG. 14 shows the SPA results of the effect of the Ga loading in the ZSM-5 zeolite. It can be observed that the unloaded zeolite exhibits the highest concentrations of ethylbenzene and xylenes in the outlet gas, whereas the addition of Ga to the zeolite favours the formation of benzene and higher hydrocarbons. A maximum overall carbon selectivity to aromatic hydrocarbons of about 73% has been determined for the 0.5 wt. % Ga-zeolite, as reported in the overview of the results in Table 1.

(24) TABLE-US-00001 TABLE 1 Summary of effect of Ga loading on ZSM-5 zeolite on the catalyst performance C.sub.2H.sub.4 conv Carbon selectivity (%) Ga loading (%)* Benzene* Toluene* Xyl** EB** C.sub.10** Total Ar*** .sup. 0% 78.9 ± 0.8  7.5 ± 1.4 43.9 ± 3.2 8.7 1.3 5.6 66.9 0.5% 87.6 ± 2.2 31.4 ± 0.8 31.5 ± 1.3 4.1 0.7 5.5 73.2 2.5% 96.7 ± 0.4 24.4 ± 1.1 31.8 ± 1.3 3.6 0.2 7.0 67.1 *From average values measured by online micro-GC analysis. **From SPA analysis taken under stable conditions (after 2 hours on stream). Xyl = xylenes (o/m/p); EB = ethylbenzene; C.sub.10 = naphthalene + 1-methyl-naphthalene + 2-methylnaphthalene. ***Total selectivity to aromatic compounds.

CONCLUSIONS

(25) The catalytic conversion of ethylene present in product gas from biomass gasification into aromatics (BTX) using bifunctional Ga-loaded ZSM-5 zeolites as catalyst under realistic gasification conditions has been analysed. The results have shown that ethylene conversion of 80-97% can be achieved. In all cases, acetylene conversion was complete. The carbon contained in ethylene and acetylene is mainly converted to benzene, toluene, ethane and methane. Ethane and methane (favourable compounds in view of bio-SNG production) are intermediate by-products of hydrodealkylation reactions. SPA analysis has revealed the formation presence of other aromatic compounds, namely xylenes, ethylbenzene, naphthalene, and naphthalene derivatives. The addition of Ga to the zeolite significantly improved both the ethylene conversion (90-97%) and the carbon selectivity to benzene. The 0.5 wt. % promoter-zeolite achieved the highest carbon selectivity to benzene (˜32%), benzene and toluene (˜65%), and total carbon selectivity to aromatics (73%). Moreover, it has been observed that the reaction temperature dramatically influences the distribution of carbon selectivity towards aromatics. Lower temperatures favour the production of ethylbenzene and xylenes, whereas benzene, naphthalene and naphthalene derivatives are promoted at higher temperatures. Based on the results, it is proposed that the formation of benzene and toluene need different active sites. The location of the active sites of ethylene conversion to toluene may be mainly on the surface of zeolite, whereas the active sites for benzene formation may be located in the pores of the zeolite. Moreover, Ga produces a partial replacement of the zeolite acid sites within the pores of the zeolite which eventually modifies the extent of the aromatic interconversion reactions (hydrodealkylation) toward benzene formation.

Example 2

(26) Beech wood (5 kg/h) was subjected to gasification with 1 kg steam per h in an indirect allothermal biomass gasifier (MILENA), which is coupled to a pre-washing unit (OLGA absorber). A trace amount of argon, which was used as tracker, was added to the gasifier gas, which was fed to the pre-washing unit with an average gas flow of about 15 Nl/min, based on dry gas. The partly cleaned gas was led via a cooler (T=5° C.), a safety filter (soxhlet filter) and glass beads to an absorber (BTX scrubber). The absorber of the BTX scrubber operated at 35° C. and ambient pressure with polymethylphenylsiloxane as a washing liquid. The loaded absorbent was stripped at 120° C. using 205-820 g/h steam. Loaded stripping gas was led to condensers via tubes heated at 120-160° C. The first condenser operated at 25-27° C. and the second condenser at 4-5° C. Liquids were collected from the first condenser, while remaining gases were led to the second condenser. Liquids were collected from the second condenser. The scrubbed gas was subjected to HDS and subsequently steam reforming to obtain a bio-SNG. The experiment was run continuously for 75 hours at ambient pressure.

(27) Collection of liquids occurred by first collecting the aqueous layer by opening a tap at the bottom of the collection flask. Collection of the aqueous layer was stopped just prior to the meniscus reached the tap. An as small as possible mixed fraction was collected and discarded, after which the organic layer was completely drained in a separate flask. A total of 1.17 kg of organic layer was collected (885 g from the first condenser at 26° C. and 285 g from the second condenser at 5° C.) over the complete duration of the experiment, of which 86.6 wt % benzene, 6.5 wt % toluene and 0.20 wt % xylene. A detailed compositional analysis of the combined organic layers is given in Table 2. The average compositions of the gas flows over the complete duration of the experiment, as determined by micro-GC, are given in Table 3.

(28) TABLE-US-00002 TABLE 2 Composition of the combined organic layers obtained from the BTX scrubber. First condenser Second condenser Compound (26° C.) (4-5° C.) Total total (g) 884.5 284.7 1169.2 total BTX (g) 824.6 266.8 1091.4 benzene (g, wt %) 753.6 (85.2 wt %) 258.9 (90.9 wt %) 1012.5 (86.6 wt %) toluene (g, wt %) 68.7 (7.77 wt %) 7.77 (2.73 wt %) 76.50 (6.54 wt %) xylene (g, wt %) 2.26 (0.26 wt %) 0.06 (0.02 wt %) 2.32 (0.20 wt %) ethylbenzene (wt %) 0.19 0.02 0.15 styrene (wt %) 1.48 0.07 1.14 cresol (wt %) 0.37 0.00 0.28 naphthalene (wt %) 0.75 0.00 0.57 further aromatic 0.82 0.39 0.72 compounds (wt %) thiophene (wt %) 0.12 0.13 0.12 water (wt %) 0.07 0.67 0.22

(29) TABLE-US-00003 TABLE 3 Composition of the gas flows (based on dry volume) Gasifier pre-washed purified Component gas .sup.[a] gas .sup.[b] gas .sup.[c] inert (vol %) .sup.[d] 3.7 4.8 5.4 CH.sub.4 (vol %) 11.3 10.9 10.5 CO (vol %) 32.0 28.9 23.7 CO.sub.2 (vol %) 25.3 25.0 27.1 C.sub.2 (vol %) .sup.[e] 3.9 3.9 3.0 H.sub.2S (ppmV) 100 151 88 COS (ppmV) 3 5 0 benzene (ppmV) 8131 4680 271 toluene (ppmV) 596 261 0 thiophene (ppmV) 20 17 0.9 tar (mg/Nm.sup.3) 17565 680 61 H.sub.2 (vol %) .sup.[f] 22.9 26.0 30.0 .sup.[a] gas emerging from the gasifier, prior to being subjected to pre-washing; .sup.[b] gas emerging from the OLGA absorber, prior to being subjected to the BTX scrubber; .sup.[c] gas emerging from the BTX scrubber; .sup.[d] Ar + N.sub.2; .sup.[e] ethane + ethylene + acetylene; .sup.[f] H.sub.2 content estimated, based on total volume of 100 vol %.

(30) Tar-like components were mainly removed in the pre-washing step, while monocyclic aromatic compounds, such as benzene, toluene and even thiophene, were largely maintained in the permanent gas stream. The BTX scrubber effectively removed the monocyclic aromatic components. The benzene concentration in the partly cleaned gas emerging from the OLGA absorber was between 4000 and 7000 ppm (vol.), which was lowered to ˜300 ppm (vol.) in the gas stream emerging from the BTX-scrubber. The average removal of benzene amounted to 95% using a steam flow of 820 g/h, which reduced to 89% and 87% at a gas flow of 410 g/h and 205 g/h respectively. 100% of the toluene was removed at all gas flows. Only trace amounts of tar-like components were obtained in the organic layers obtained in the first and second condensers.

(31) The composition of the gas was analysed prior to and after the condensers, and the removal percentages obtained during the BTX scrubbing of a variety of compounds are given in Table 4. The aromatic compounds benzene, toluene, xylene and thiophene (and its derivatives) were effectively removed during the BTX scrubbing, wherein generally the highest steam flow provided the highest removal. At the same time, permanent gases such as C.sub.1-C.sub.3 hydrocarbons (alkanes and alkenes) are effectively retained in the gas stream. Especially methane, CO and CO.sub.2 are completely retained. Any transport thereof to the stripping gas is cancelled when the permanent gases are recycled to the entrance of the BTX scrubber. The content of the permanent gases in the fuel gas (at the entrance of the BTX scrubber) and in the permanent gas stream after stripping (downstream of the second condenser) is given in Table 5. In view of its very small volume, nitrogen gas was added to the permanent gas stream to enable measure of its contents (tracer). The amount of these permanent gases that were transported to the stripping gas is also included in Table 5. The permanent gas stream further contained 4.6 vol % benzene (based on the permanent gas stream without added nitrogen).

(32) TABLE-US-00004 TABLE 4 Removal percentages for certain compounds (15 NL/min inlet gas) steam flow during stripping Compound 205 g/h 410 g/h 820 g/h Aromatic compounds benzene 87% 89% 95% toluene 93% 92% 100%  xylene 100%  100%  100%  Sulphur components thiophene 94% 91% 96% 2-methyl-thiophene .sup.[b] 100%  100%  100%  3-methyl-thiophene .sup.[b] 100%  100%  100%  COS  6%  0% 14% methyl mercaptan 67% 68% 60% ethyl mercaptan 76% 78% 77% [a] nd = not determined .sup.[b] no methyl-thiophenes were detected after the BTX-scrub.

(33) TABLE-US-00005 TABLE 5 Experimentally determined transport values of permanent gases Permanent gases In fuel gas In strip gas transported methane (vol %) 11.00 0.13 1.18 ethane (vol %) 0.20 0.0094 4.70 ethene (vol %) 3.10 0.168 5.43 ethyne (vol %) 0.155 0.0185 11.94 CO.sub.2 (vol %) 25.00 0.72 2.88 CO (vol %) 28.00 0.07 0.25 H.sub.2 (vol %) 26.00 0.00 0.00 H.sub.2S (ppmV) 150.00 63 42.00 COS (ppmV) 5.00 0.00 0.00

(34) The results in Table 4 show that the BTX-scrubber, i.e. step (b) of the process according to the invention, effectively removes a fraction of aromatic compounds, which contains almost exclusively BTX. The BTX fraction can be used as deemed fit, e.g. marketed as bio-based BTX or the like, while the energy gas is sufficiently purified from tar-like components and aromatic components by virtue of the combined pre-washing and BTX-scrubber such that conversion into bio-SNG (or other products) is readily accomplished.

Example 3

(35) Refuse-derived fuel (RDF) (3.9 kg/h) was subjected to gasification with 2 kg steam per h in an indirect allothermal biomass gasifier (MILENA), which is coupled to a pre-washing unit (OLGA absorber). Nitrogen gas was added to the steam flow to maintain sufficient gas flow throughout the system. The gasifier gas was fed to the pre-washing unit with an average gas flow of about 15 Nl/min, based on dry gas. The partly cleaned gas was led via a cooler (T=5° C.), a safety filter (glass beads) and a pre-washing unit to remove tars (washing liquid=polymethylphenylsiloxane; T=80° C., ambient pressure) to an absorber (BTX scrubber). The absorber of the BTX scrubber operated at 35° C. and ambient pressure with polymethylphenylsiloxane as a washing liquid. The loaded absorbent was stripped at 160° C. using 0.25 m.sup.3/h steam. Loaded stripping gas was led to condensers via tubes heated at 160° C. The first condenser operated at 25-27° C. and the second condenser at 4-5° C. Liquids were collected from the first condenser, while remaining gases were led to the second condenser. Liquids were collected from the second condenser. A fraction of monocyclic aromatic compounds comprising benzene, toluene and xylene was collected. The experiment was run continuously for 3.5 hours at ambient pressure. The average compositions of the gas flows over the complete duration of the experiment, as determined by micro-GC, are given in Table 6.

(36) TABLE-US-00006 TABLE 6 Composition of the gas flows Gasifier pre-washed purified Component gas .sup.[a] gas .sup.[b] gas .sup.[c] inert (vol %) .sup.[d] 51.7 53.8 53.4 CH.sub.4 (vol %) 7.8 7.7 7.5 CO (vol %) 9.0 9.0 9.0 CO.sub.2 (vol %) 11.4 12.2 11.2 C.sub.2 (vol %) .sup.[e] 7.4 7.8 7.2 H.sub.2S (ppmV) 524 633 449 COS (ppmV) 9 16 0 benzene (ppmV) 9853 9135 302 toluene (ppmV) 1490 1269 35 thiophene (ppmV) 43 29 nd .sup.[f] tar (mg/Nm.sup.3) 38373 809 61 H.sub.2 (vol %) .sup.[g] 8.4 8.4 8.4 .sup.[a] gas emerging from the gasifier, prior to being subjected to pre-washing; .sup.[b] gas emerging from the OLGA absorber, prior to being subjected to the BTX scrubber; .sup.[c] gas emerging from the BTX scrubber; .sup.[d] Ar + N.sub.2; .sup.[e] ethane + ethylene + acetylene; .sup.[f] not determined; .sup.[g] H.sub.2 content estimated, based on total volume of 100 vol %.

Example 4

(37) The Ga-ZSM-5 aromatization catalyst with a loading of about 2.5 wt %, synthesized and tested in Example 1, was re-used for a duration test for the determination of the long-term performance of the catalyst. All process steps and conditions (temperature, GHSV, gasification conditions) were the same as in Example 1. The catalyst was in operation for approximately 25 hours at a fixed temperature of 500° C. The catalyst performed stably over the entire run, with only a slight overall decrease in ethylene conversion. Ethylene conversion was about 95-97%. The carbon selectivity of ethylene and acetylene to aromatics was stable over the 25 hours of the duration test. The catalyst showed carbon selectivity to benzene of about 20%, and of about 30% to toluene. The carbon selectivity to xylenes was in comparison significantly lower than that of benzene and toluene (about 5% in total, with slightly higher selectivity to m- and p-xylene than to o-xylene). The catalyst showed negligible selectivity to ethylbenzene. These results are depicted in FIG. 15.

(38) The inlet and outlet flows of the reactor were determined for ethylene, acetylene, ethane, benzene, toluene, ethyl-benzene and xylene. Gas concentrations were measured during the entire run online by micro-GC (as in Example 1), and neon was added as tracer gas for the determination of molar balances around the reactor. All inlet and outlet flows remained stable for the entire run. The results are depicted in Table 7.

(39) TABLE-US-00007 TABLE 7 Inlet and outlet concentrations of selected gas compounds (operation at 500° C.) Inlet gas concentration Outlet gas concentration Compound (on dry basis) (on dry basis) ethylene  .sup. 3.89 ± 0.04 vol. %  .sup. 0.17 ± 0.02 vol. % acetylene  .sup. 0.37 ± 0.007 vol. % <0.001 vol. % * ethane  .sup. 0.24 ± 0.005 vol. %  .sup. 1.38 ± 0.08 vol. % methane  .sup. 12.4 ± 0.13 vol. %  .sup. 13.4 ± 0.12 vol. % benzene 7554.5 ± 289.9 ppmv 10607.7 ± 180.8 ppmv  toluene 554.7 ± 27.9 ppmv 4294.0 ± 118.0 ppmv ethylbenzene  19.1 ± 15.1 ppmv 31.7 ± 5.4 ppmv m/p-xylene 14.6 ± 0.7 ppmv 286.3 ± 19.7 ppmv o-xylene  6.4 ± 0.6 ppmv 144.1 ± 11.6 ppmv * Detection limit of micro-GC

(40) Ethylene was converted down to 0.2 vol. %, whereas acetylene was converted below detection limits. The concentration of ethane was slightly increased in the reactor, which was also the case for methane. Since ethane and methane are desirable components of (bio-)SNG, the process according to the invention was not only able to convert ethylene into useful monocyclic aromatic compounds, but also the yield of bio-SNG was increased. Benzene concentration was increased with 30%, while toluene outlet flow was 7 times as high as inlet flow. For the xylene isomers, the outlet flow was about 20 times higher as the inlet flows. On the contrary, the carbon selectivity of the catalyst to ethylbenzene was negligible. Carbon selectivity to xylenes and ethylbenzene was significantly favoured when operating at lower temperatures (data not shown).

(41) Further, the ZSM-5 support was found to completely capture H.sub.2S and COS. The breakthrough of H.sub.2S took place almost 6 hours after the start of the run. However, the zeolite still retained certain capacity for sulphur capture, which gradually decreases over time. After 25 hours operation at 500° C., the zeolite still captured approximately half of the H.sub.2S. The N.sub.2 flush applied to the catalyst after 25 hours operation seemed to partially regenerate the zeolite. The zeolite support was also able to capture thiophene and mercaptan derivatives (not shown), which is beneficial for the purity of the liquid product containing the monocyclic aromatic compounds.