feed, and H 2 product pressures are high enough to feed to downstream hydrotreaters without need for any additional compression. It also tends to have higher H 2 purities in the ROG, with concentrations up to 70-80 mol%. For each hydrotreater and hydrocracker case, perme- ate purities from 95-99.9% were considered (see Table 1 ). When there is flexibility in the purity of produced H2 , membrane separation is incredibly beneficial. On average, a 99.9% purity H 2 stream requires five times the mem - brane Capex investment of a 95% purity stream. These H2 streams leave the membrane systems at a lower pressure than the feed to be sent either directly to any downstream unit operating at lower pressures or to recycle compressors. Ancillary equipment, including valves, piping, and pressure control, would add ~50% to the module cost. The high-pressure ROG streams produce a purified H2 stream that is still at high pressure. These streams could easily be recycled to lower-pressure systems without any recompression required. For a stream like the naphtha hydrotreater ROG, the H 2 permeate would only be between 7-17 bar. It would almost certainly need to be recom - pressed to either be recycled back to the source unit or uti - lised elsewhere in the refinery. It is important to keep in mind that another benefit of membrane separation is that the retentate stream (the H2 poor stream) leaves the membrane system at a near-feed pressure, allowing for a variety of different recovery and recycle opportunities. Assuming there is available H2 recycle compressor capac- ity, the costs of recompression to get the H 2 stream back up to feed pressure can be considered. For this analysis, we will look at the H 2 recovery over the life of a membrane unit (five years) and assume a unit uptime of 95% (to be conservative) with total H 2 production. The cost of H2 sep- aration will then be compared back to the cost of producing new H 2 via a grey H 2 steam methane reformer (SMR) pro- cess (with an estimated price per kg of $1.80) to find the breakeven point:
Where BEP days = breakeven point (days) C mem = Capex cost of membrane modules ($) C anc = Capex cost of ancillary equipment ($) C H2gen = Cost per kg fresh H2 generated ($/kg) R H₂rec = Rate of H₂ permeate production (kg/hr) P comp = Compression power required (kW) E cost = Electricity cost ($/kWh)
With the processing units shown in Table 2 , all options show a positive ROI but are highly variable depending on stream conditions and required purities. This demonstrates the criticality of designing the membrane system only to the required purity rather than automatically matching the 99.9% purity attainable from a pressure swing adsorption (PSA) system. The total H₂ recovery values vs the amount spent in mem - brane system Capex and compression Opex are significant, with the naphtha hydrotreater 95% purity case recovering more than $121M in H 2 value with only a $3.2M Capex investment and $1.8M in compression costs over the mem- brane life. Figure 1 is a visual comparison of each case vs the cost of new H₂ from a grey SMR H2 plant, including the Opex cost for compression back to feed pressures. Isomerisation unit, catalytic reformer, FCC The isomerisation unit typically has much lower ROG flow rates than the hydrotreaters and hydrocracker to upgrade low-quality naphtha into higher-quality gasoline blending components. The catalytic reformer is one of the few units in the refinery that produces H2 as a side product to its main reactions. It frequently supplements the SMR in H 2 production for the refinery. H2 from this ROG stream is typ- ically captured via PSA, and this stream will be reviewed for future comparison to PSA separation costs. H 2 recovery from the fluid catalytic cracker (FCC) was not assessed. A combination of it is low operating pres - sure (2 bar) and low H₂ purity (<10%) make it less ideal for membrane separation processes. The stream would require compression to at least 15 bar before feeding to the mem - brane module, and the H₂ permeate stream would require additional compression to be useful. While the separation
Stream
Recovery (kg/hr H 2)
Compressor
Module cost + ancillary +
BEP
H₂ value
Total Capex
Total Five-year
power
(months) over five
Opex over ROI
required (kW) compression ($/kg)
years
five years
GO hydrotreater 95% purity GO hydrotreater 99% purity GO hydrotreater 99.9% purity Hydrocracker 95% purity Hydrocracker 99% purity Hydrocracker 99.9% purity Diesel hydrotreater 95% purity Diesel hydrotreater 99% purity Diesel hydrotreater 99.9% purity Naphtha hydrotreater 95% purity Naphtha hydrotreater 99% purity
1,262 1,260 1,201 405.3 404.9 384.4 1,021 1,020 965.9 1,621 1,626
751.5 430.7
$0.07 $0.33 $0.63 $0.27 $0.30 $0.68 $0.09 $0.17 $0.55 $0.17 $0.62 $0.96
0.75 9.78
$94.5M $94.3M $89.9M $30.4M $30.3M $28.8M $76.5M $76.4M $72.3M $121.4M $121.8M $115.8M
$1.2M
$2.5M $1.4M $3.9M $1.0M $1.3M $2.0M $1.4M $3.3M $2.7M $7.9M $1.8M $4.7M
2,454%
$15.7M $27.8M
450% 184% 572% 505% 166%
1,156.6
18.69
311.5 387.9 588.2 426.4 981.7 801.7 2388 535.3
6.86 7.40
$3.5M $3.7M $8.9M $2.3M $4.1M
19.13
1.76 3.22
1,967%
940% 229% 985% 189%
16.02
$19.3M
1.65
$3.2M
19.42 29.72
$40.3M $57.1M
Naphtha hydrotreater 99.9% purity 1,546
1,414.7
87%
Table 2 Hydrotreater and hydrocracker economic analysis with recompression to feed pressures
16
PTQ Q3 2024
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