100
1.00
8
4
Benzene
80
2
0.80
1
60
Slurry
0.60
40
0.40
20
No. stages
0.20
0
0
1
2
3
4
5
6
Volume stripping/Volume emulsion gas
60.0 70.0 0 10.0 20.0 30.0 40.0 50.0 Time (seconds) 0.00
100.0 80.0 90.0
Figure 2 Stripping at 500ºC
Figure 3 Expected performace from ideal stages
stages should give excellent performance and can be installed in most applications. A graphical representation of this is shown in Figure 2 . It suggests that ideally eight stages with a steam volume of two times the emulsion phase should be adequate. In gasoil applications, the 2 lb/1,000 lb of catalyst cir- culated criteria works for disc and doughnut strippers operating within their design guidelines. These were 2-3 lb steam/1000 lb catalyst circulated, a catalyst flux rate of Reaction mix sampling can be used to assess the stripper performance and the amount of cracking occurring downstream of the riser reactor 600-900 lb catalyst/ (ft² minute), and a catalyst residence time of 60-90 seconds. Stripping tests in the laboratory for benzene and slurry at 500ºC (932ºF) are shown in Figure 3 . Both streams are aro- matic and represent the hardest to desorb hydrocarbons. Benzene would be essentially inert in the stripper, while the slurry would thermally crack at the long residence times of the catalyst stripper. Having a residence time longer than 90 seconds might give a lower hydrogen in coke but would add to the wet gas compressor load. The dry gas produced
is high in hydrogen content and lowers the molecular weight of the gas going to the wet gas compressor. Higher reactor temperatures increase the desorption rate and reduce the time needed to effectively strip the catalyst. Reaction mix sampling can be used to assess the stripper performance and the amount of cracking occur- ring downstream of the riser reactor. Cracking in the strip- per will increase the delta coke and raise the regenerator temperature. The yield and operating changes used to calculate the typical impact of a stripper revamp are shown in Table 2 . A higher regenerator temperature lowers the catalyst/oil ratio and conversion. Gasoline declines, but decant oil (DO) and dry gas both increase. The slightly lower coke yield is due to the higher heat content (hydrogen) of the coke. These changes are significant but not large enough to change out a disc and doughnut stripper with a larger diameter to lower the flux rate. Most strippers operating with more than 7 wt % hydrogen in coke are well above the design catalyst flux rates. The pinch points in the disc and doughnut design cause high localised catalyst velocities that impede the rise of small bubbles. They either coalesce into larger bubbles or get swept down the stripper to the catalyst exit. Individual bubbles rise according to their size, as shown in Table 3 . Higher values are observed with clus- ters of bubbles. Replacing the shell of the stripper is very expensive and results in a project having a three- to four- year payback. These payback times seldom get funding.
Overall efficiency
Stage efficiency
Number of stages
1
2
3
4
5
6
7
8
9
10
10 20 30 40 50 60 70
10 20 30 40 50 60 70
19 36 51 64 75 84 91
27.1 48.8 65.7 78.4 87.5 93.6 97.3
34.4 59.0 76.0 87.0 93.8 97.4 99.2
41.0 67.2 83.2 92.2 96.9 98.9 99.8
46.9 73.8 88.2 95.3 98.5 99.5 99.9
52.2 79.0 91.7 97.2 99.2 99.8 100
57.0 83.2 94.2 98.3 99.6 99.9 100
61.3 86.6 95.9 99.0 99.8 100 100
65.2 89.3 97.1 99.4 99.9 100 100
Table 1
74
PTQ Q1 2025
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