My experience as a Process Engineeer is that I would supply operating guidelines. Usually a high level overview of the process. This is then turned over to operations to develop detailed operating procedures.
If I remember this correctly...
You conduct the UT or RT to find the minimum wall thickness remaining in the pipe. Then using the minimum wall thickness remaining calculate the maximum allowable operating pressure.
Did you not have a plant design course as part of your curricilum? I am sure there are many books available that can help you.
The most knowledge you are going to gain is through work experience.
To MJCronin:
I think you should go back and take a read of what I wrote.
"I am not sure of the exact reason but it is pretty typical to only specify half vacuum for vessels that are only steamed out for maintenance. That is what is specified in our design manuals."
"Our" being the EPC I work...
I am not sure of the exact reason but it is pretty typical to only specify half vacuum for vessels that are only steamed out for maintenance. That is what is specified in our design manuals. Maintenanace procedures would dictate that the vessel is not to be blocked in during steam out. My guess...
No you don't have to align the piping design pressure with the casing MAWP. I usually set the downstream piping design pressure as the maximum discharge pressure - which is shut off head plus maximum suction pressure. However, keep in mind that if you ever put a bigger impeller in the pump your...
So you have a sample loop that comes of your pump discharge is regulated down to 50 psi and then returns to the suction side. So to see greater than 50 psi you would have to have your regulator fail, the pump suction valve closed and the valve closed at the deaerator closed as well?? A sketch...
I have seen it done both ways. However, my experience has been that the design pressure as shown in the line list is maxiumum pressure that the line can actually see. For example, if the line is connected to a vessel that is protected by a PSV then the maximum pressure the line can see is the...
You need to look at the required Cv at your normal and minimum flow operating conditions as well. At the maximum the valve appears to be suitable. Check the required Cv and valve opening at the minimum and normal condition. That is where you might run into problems.
Back to the sump is the safer method. By going back into the suction line you are essentially creating a recirculation loop - fluid from discharge goes thru PSV into suction line and back through pump again. So, system is protected from overpressure but now you are starting to heat up the...
It depends on the Cv of the valve. If you have the valve make and model you can contact the vendor and have them check rate the valve for the new operating conditions. If it is a Fisher valve you can download software to do the check rate yourself.
If there is any valve in the 150# line that can be closed to create a blocked discharge situation you have the potential to overpressure the 150# system. If you can't CSO the valve then you would need a PSV to protect the line.
Your contractor is proposing an ESDV. I assume to trip closed on...
This is why I would like to see a sketch. You will have to carry the 300# rating up to the last valve in the system. Any valves in the 150# system will have to be CSO or better Locked Open or you have to install a PSV in the 150# system.
You have 3 posts on here regarding gas sweetening. Wouldn't some of these questions be better directed at the Process Engineer that I hope is working with you?
